Process for treating a hydrocarbon-containing feed

ABSTRACT

A process for treating a hydrocarbon-containing feedstock is provided in which a hydrocarbon-containing feed comprising at least 20 wt. % of heavy hydrocarbons is mixed with hydrogen and at least one catalyst to produce a hydrocarbon-containing product. The hydrocarbon-containing feedstock, the catalyst(s), and the hydrogen are provided to a mixing zone and blended in the mixing zone at a temperature of from 375° C. to 500° C. A vapor comprised of hydrocarbons that are vaporizable at the temperature and pressure within the mixing zone is separated from the mixing zone, and, apart from the mixing zone, the vapor is condensed to produce a liquid hydrocarbon-containing product. The hydrocarbon-containing feedstock is continuously or intermittently provided to the mixing zone at a rate of at least 350 kg/hr per m 3  of the mixture volume in the mixing zone.

CROSS REFERENCE TO RELATED APPLICATION

The present application claims the benefit of priority from U.S.Provisional Patent Application Ser. No. 61/297,082 filed Jan. 21, 2010.

FIELD OF THE INVENTION

The present invention is directed to a process for treating ahydrocarbon-containing feedstock.

BACKGROUND OF THE INVENTION

Increasingly, resources such as heavy crude oils, bitumen, tar sands,shale oils, and hydrocarbons derived from liquefying coal are beingutilized as hydrocarbon sources due to decreasing availability of easilyaccessed light sweet crude oil reservoirs. These resources aredisadvantaged relative to light sweet crude oils, containing significantamounts of heavy hydrocarbon fractions such as residue and asphaltenes,and often containing significant amounts of sulfur, nitrogen, metals,and/or naphthenic acids. The disadvantaged crudes typically require aconsiderable amount of upgrading, for example by cracking and byhydrotreating, in order to obtain more valuable hydrocarbon products.Upgrading by cracking, either thermal cracking, hydrocracking and/orcatalytic cracking, is also effective to partially convert heavyhydrocarbon fractions such as atmospheric or vacuum residues derivedfrom refining a crude oil or hydrocarbons derived from liquefying coalinto lighter, more valuable hydrocarbons.

Numerous processes have been developed to crack and treat disadvantagedcrude oils and heavy hydrocarbon fractions to recover lighterhydrocarbons and to reduce metals, sulfur, nitrogen, and acidity of thehydrocarbon-containing material. For example, a hydrocarbon-containingfeedstock may be cracked and hydrotreated by passing thehydrocarbon-containing feedstock over a catalyst located in a fixed bedcatalyst reactor in the presence of hydrogen at a temperature effectiveto crack heavy hydrocarbons in the feedstock and/or to reduce the sulfurcontent, nitrogen content, metals content, and/or the acidity of thefeedstock. Another commonly used method to crack and/or hydrotreat ahydrocarbon-containing feedstock is to disperse a catalyst in thefeedstock and pass the feedstock and catalyst together with hydrogenthrough a slurry-bed, or fluid-bed, reactor operated at a temperatureeffective to crack heavy hydrocarbons in the feedstock and/or to reducethe sulfur content, nitrogen content, metals content, and/or the acidityof the feedstock. Examples of such slurry-bed or fluid-bed reactorsinclude ebullating-bed reactors, plug-flow reactors, and bubble-columnreactors.

Coke formation, however, is a particular problem in processes forcracking a hydrocarbon-containing feedstock having a relatively largeamount of heavy hydrocarbons such as residue and asphaltenes.Substantial amounts of coke are formed in the current processes forcracking heavy hydrocarbon-containing feedstocks, limiting the yield oflighter molecular weight hydrocarbons that can be recovered anddecreasing the efficiency of the cracking process by limiting the extentof hydrocarbon conversion that can be effected per cracking step in theprocess, for example, by deactivating the catalysts used in the process.

Cracking heavy hydrocarbons involves breaking bonds of the hydrocarbons,particularly carbon-carbon bonds, thereby forming two hydrocarbonradicals for each carbon-carbon bond that is cracked in a hydrocarbonmolecule. Numerous reaction paths are available to the crackedhydrocarbon radicals, the most important being: 1) reaction with ahydrogen donor to form a stable hydrocarbon molecule that is smaller interms of molecular weight than the original hydrocarbon from which itwas derived; and 2) reaction with another hydrocarbon or anotherhydrocarbon radical to form a hydrocarbon molecule larger in terms ofmolecular weight than the cracked hydrocarbon radical—a process calledannealation. The first reaction is desired, it produces hydrocarbons oflower molecular weight than the heavy hydrocarbons contained in thefeedstock—and preferably produces naphtha, distillate, or gas oilhydrocarbons. The second reaction is undesired and leads to theproduction of coke as the reactive hydrocarbon radical combines withanother hydrocarbon or hydrocarbon radical. Furthermore, the secondreaction is autocatalytic since the cracked hydrocarbon radicals arereactive with the growing coke particles. Hydrocarbon-containingfeedstocks having a relatively high concentration of heavy hydrocarbonmolecules therein are particularly susceptible to coking due to thepresence of a large quantity of high molecular weight hydrocarbons inthe feedstock with which cracked hydrocarbon radicals may combine toform proto-coke or coke. As a result, cracking processes of heavyhydrocarbon-containing feedstocks have been limited by coke formationinduced by the cracking reaction itself.

Processes that utilize fixed bed catalysts to crack a heavyhydrocarbon-containing material suffer significantly from catalyst agingdue to coke deposition on the catalyst over time. As noted above, cokeand proto-coke formation occurs in cracking a hydrocarbon-containingmaterial, and is particularly problematic when thehydrocarbon-containing material is a heavy hydrocarbon-containingmaterial, for example, containing at least 20 wt. % pitch, residue, orasphaltenes. The coke that is formed in the cracking process deposits onthe catalyst progressively over time, plugging the catalyst pores andcovering the surface of the catalyst. The coked catalyst loses itscatalytic activity and, ultimately, must be replaced. Furthermore, thecracking process must be conducted at relatively low crackingtemperatures to prevent rapid deactivation of the catalyst byannealation leading to coke deposition.

Slurry catalyst processes have been utilized to address the problem ofcatalyst aging by coke deposition in the course of cracking ahydrocarbon-containing feedstock. Slurry catalyst particles are selectedto be dispersible in the hydrocarbon-containing feedstock or invaporized hydrocarbon-containing feedstock so the slurry catalystscirculate with the hydrocarbon-containing feedstock in the course ofcracking the feedstock. The feedstock and the catalyst move togetherthrough the cracking reactor and are separated upon exiting the crackingreactor. Coke formed during the cracking reaction is separated from thefeedstock, and any coke deposited on the catalyst may be removed fromthe catalyst by regenerating the catalyst. The regenerated catalyst maythen be recirculated with fresh hydrocarbon-containing feedstock throughthe cracking reactor. The process, therefore, is not affected bycatalyst aging since fresh catalyst may be continually added into thecracking reactor, and catalyst upon which coke has been deposited may becontinually regenerated.

Other slurry catalysts have been used in slurry cracking processes forthe purpose of seeding the formation of coke. Very small particle slurrycatalysts may be dispersed in a hydrocarbon-containing feedstock for thepurpose of providing a plethora of small sites upon which coke maydeposit in the course of the cracking process. This inhibits theformation of large coke particles since the coke may be dispersedthroughout the hydrocarbon-containing feedstock on the small catalystparticles.

While slurry catalyst processes provide an improvement over fixed-bedcatalysis of heavy hydrocarbon feedstocks, coking remains a problem.Generally, the upper limit of recovery of hydrocarbons from a heavyhydrocarbon cracking process is around 70%, where the non-recoverablehydrocarbons are converted into coke and gas.

WO 2008/141830 and WO 2008/141831 provide a process and system forhydroconversion of heavy oils utilizing a solid accumulation reactor. Ahydrogenation catalyst is dispersed in a slurry in a reactor capable ofoperating stably in the presence of solids deriving from and generatedby a heavy oil. Heavy oil is hydroconverted to produce a lighterhydrocarbon product by reaction of the heavy oil with hydrogen and thecatalyst at temperatures effective to convert the heavy oil. Product maybe vaporized in the reactor and stripped from the slurry to be capturedas a vapor exiting the reactor, or a liquid product may be separatedfrom the reactor, where a vapor product may be separated from the liquidproduct separated from the reactor. Solids including coke and metalsproduced by the hydroconversion accumulate in the reactor and areremoved from the reactor by continuous flushing in proportion to theamount of solids generated once a pre-established minimum accumulationlevel is reached in the reactor. Large amounts of solids including coke,sulfided metals, and insoluble asphaltenes are generated in the processof producing the vapor product. As a result, the rate at which the heavyoil may be hydroconverted is quite slow, ranging from 50 to 300 kg/h m³of reaction volume.

The slow rate and the large quantities of solids produced by the processdisclosed in WO 2008/141830 and WO 2008/141831 limits the commercialusefulness of the process. Large scale commercial facilities forupgrading heavy crude oils must be capable of upgrading large quantitiesof oil rapidly—typically on the order of 100,000 barrels per day.Therefore, due to the slow rate of the process disclosed in WO2008/141830 and WO 2008/141831, a very large reactor having a largevolume capacity would be required to upgrade a heavy oil on acommercially efficient scale using the process. Such reactors areextremely capital intensive, either prohibiting or limiting theapplication of the process due to the expense of building a commerciallyeffective reactor.

Improved processes for cracking heavy hydrocarbon-containing feedstocksto produce a lighter hydrocarbon-containing crude product are desirable,particularly in which coke formation is significantly reduced oreliminated and the rate of hydroconversion is greatly increased.

SUMMARY OF THE INVENTION

A process for cracking a hydrocarbon-containing feedstock, comprising:

continuously or intermittently providing hydrogen to a mixing zone;providing a metal-containing catalyst to the mixing zone;selecting a hydrocarbon-containing feedstock containing at least 20 wt.% hydrocarbons having a boiling point of greater than 538° C. asdetermined in accordance with ASTM Method D5307;continuously or intermittently providing the hydrocarbon-containingfeedstock to the mixing zone at a selected rate and blending thehydrogen, the hydrocarbon-containing feedstock, and the catalyst in themixing zone at a temperature of from 375° C. to 500° C. and at a totalpressure of from 6.9 MPa to 27.5 MPa to produce:

-   -   a) a vapor comprised of hydrocarbons that are vaporizable at the        temperature and the pressure within the mixing zone, and    -   b) a hydrocarbon-depleted feed residuum comprising hydrocarbons        that are liquid at the temperature and pressure within the        mixing zone,    -   where the combined volume of the hydrocarbon-depleted feed        residuum, the catalyst, and the hydrocarbon-containing feedstock        in the mixing zone define a mixture volume in the mixing zone,        wherein the rate at which the hydrocarbon-containing feedstock        is provided to the mixing zone is selected to be at least 350        kg/hr m³ of the mixture volume in the mixing zone; and        separating at least a portion of the vapor from the mixing zone        while retaining in the mixing zone at least a portion of the        hydrocarbon-depleted feed residuum.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic of a system useful for practicing the process ofthe present invention.

FIG. 2 is a schematic of a system useful for practicing the process ofthe present invention including a reactor having three zones.

FIG. 3 is a plot of hydrocracking reaction rates relative to hydrogensulfide present in the reaction.

DETAILED DESCRIPTION OF THE INVENTION

The present invention is directed to a process for cracking ahydrocarbon-containing feedstock containing at least 20 wt. % heavyhydrocarbons at a relatively rapid rate while producing little, if any,coke. A heavy hydrocarbon feedstock may be provided for cracking at aflow rate of at least 350 kilograms(kg)/hour(h) per m³ (cubic meter) ofreaction volume.

The process of the present invention may be conducted at such highrates, in part, because the 1) process is effective to crack a heavyhydrocarbon-containing feedstock while producing little coke or othertoluene-insoluble carbonaceous solids; and 2) the process may beconducted at relatively high temperatures, for example at least 450° C.,since the process inhibits formation of coke normally formed insubstantial quantities at such temperatures. Although not intending thepresent invention to be limited thereby, it is believed that theproduction of coke is inhibited in the process of the invention, and theprocess may be conducted at relatively high temperatures, in part,because the metal-containing catalyst that may be utilized in theprocess is particularly effective at selectively directing reactionsoccurring in the cracking process to avoid and/or inhibit cokeformation, and in part, since hydrogen sulfide, when utilized in theprocess, further catalyzes the cracking reactions to increase the rateof the reactions and inhibits annealation of cracked hydrocarbons, alsodirecting reactions occurring in the cracking process to avoid and/orinhibit coke formation.

Although not intending the present invention to be limited thereby, withrespect to the one or more metal-containing catalysts that may beutilized in the process, it is believed that the metal-containingcatalyst(s) are highly effective for use in cracking a heavyhydrocarbon-containing material at a high rate due, at least in part, 1)to the ability of the catalyst(s) to donate or share electrons withhydrocarbons (i.e. to assist in reducing the hydrocarbon when thehydrocarbon is cracked so the hydrocarbon forms a hydrocarbon radicalanion); and 2) the surface area of the catalyst available to interactwith hydrocarbons and/or hydrocarbon radicals in the absence of anyporous alumina, alumina-silica, or silica based carrier or support. Theone or more metal-containing catalysts that may be utilized in theprocess of the present invention have little or no acidity, andpreferably are Lewis bases.

It is believed that the hydrocarbons of a hydrocarbon-containingfeedstock are cracked in the process of the present invention by a Lewisbase mediated reaction, wherein the metal-containing catalystfacilitates a reduction at the site of the hydrocarbon where thehydrocarbon is cracked, forming two hydrocarbon radical anions from aninitial hydrocarbon compound. Hydrocarbon radical anions are most stablewhen present on a primary carbon atom, therefore, formation of primaryhydrocarbon radical anions may be energetically favored when ahydrocarbon is cracked in accordance with the process of the presentinvention, or the cracked hydrocarbon may rearrange to form the moreenergetically favored primary radical anion. Should the primary radicalanion react with another hydrocarbon to form a larger hydrocarbon, thereaction will result in the formation of a secondary carbon-carbon bondthat is susceptible to being cracked again. However, since hydrocarbonradical anions are relatively stable they are likely to be hydrogenatedby hydrogen present in the reaction mixture rather than react withanother hydrocarbon in an annealtion reaction, and significanthydrocarbon radical anion-hydrocarbon reactions are unlikely. As aresult, little coke is formed by agglomeration of cracked hydrocarbons.

Conventional hydrocracking catalysts utilize an active hydrogenationmetal, for example a Group VIII metal such as nickel, on a supporthaving Lewis acid properties, for example, silica, alumina-silica, oralumina supports. The acidic support catalyzes cracking hydrocarbons andthe active hydrogenation metal catalyzes hydrogenation of the crackedhydrocarbon radicals. It is believed that cracking heavy hydrocarbons inthe presence of a catalyst having significant acidity results in theformation of cracked hydrocarbon radical cations rather than hydrocarbonradical anions. Hydrocarbon radical cations are most stable when presenton a tertiary carbon atom, therefore, cracking may be energeticallydirected to the formation of tertiary hydrocarbon radical cations, or,most likely, the cracked hydrocarbon may rearrange to form the moreenergetically favored tertiary radical cation. Hydrocarbon radicalcations are unstable relative to hydrocarbon radical anions, and mayreact rapidly with other hydrocarbons. Should the tertiary radicalcation react with another hydrocarbon to form a larger hydrocarbon, thereaction may result in the formation of a carbon-carbon bond that is notsusceptible to being cracked again. As a result, coke is formed byagglomeration of the cracked hydrocarbons in a cracking processutilizing a conventional cracking catalyst having an acidic support orcarrier.

Although the process of the invention is not to be limited thereby, itis also believed that hydrogen sulfide, when present in significantquantities, acts as a further catalyst in the cracking of hydrocarbonsin the hydrocarbon-containing feedstock. Hydrogen sulfide, insignificant quantities, inhibits the formation of coke in the process ofcracking hydrocarbons in the hydrocarbon-containing feedstock in thepresence of hydrogen and a Lewis basic metal-containing catalyst and inthe absence of a catalyst having significant acidity. It is believedthat hydrogen sulfide, in absence of significant catalytic acidity,lowers the activation energy required to crack hydrocarbons in ahydrocarbon-containing feedstock, thereby increasing the rate of thereaction. The rate of the process, in particular the rate that thehydrocarbon-containing feedstock may be provided for cracking andcracked, hydrogenated product may be produced, therefore, may be greatlyincreased with the use of significant quantities of hydrogen sulfide inthe process. For example, the rate of a cracking process may beincreased by at least 1.5 times, or by at least 2 times, the rate of theprocess in the absence of significant quantities of hydrogen sulfide.

Hydrogen sulfide and hydrogen each may act as an atomic hydrogen donorto a cracked hydrocarbon radical anion to produce a stable hydrocarbonhaving a smaller molecular weight than the hydrocarbon from which thehydrocarbon radical was derived. Hydrogen, however, may only act todonate a hydrogen atom to a cracked hydrocarbon radical at or near ametal-containing catalyst surface. Hydrogen sulfide, however, may act todonate a hydrogen atom to a cracked hydrocarbon radical significantlyfurther from the metal-containing catalyst surface, and, after donationof a hydrogen atom, may accept a hydrogen atom from hydrogen near thesurface of the catalyst. The hydrogen sulfide, therefore, may act as anatomic hydrogen shuttle to provide hydrogen atoms to a crackedhydrocarbon radicals at a distance from the metal-containing catalyst.

Furthermore, the thiol group remaining after hydrogen sulfide hasprovided a hydrogen atom to a cracked hydrocarbon radical may beprovided to another hydrocarbon radical, thereby forming a meta-stablethiol-containing hydrocarbon. This may be described chemically asfollows:

R—C≡C—R+heat+catalyst⇄R—C

+

C—R  1.

-   -   (catalyst=basic metal-containing catalyst)

R—C

+H₂S⇄R—CH+

SH  2.

C—R+

SH⇄R—C—SH  3.

R—C—SH+H₂⇄RCH+H₂S  4.

The thiol of the meta-stable thiol-containing hydrocarbon may bereplaced by a hydrogen atom from either another hydrogen sulfidemolecule or hydrogen, or may react intramolecularly to form a thiophenecompound as a hydrocarbon-containing product.

It is believed, therefore, that hydrogen sulfide may increase the rateof the reaction 1) by lowering the activation energy of the hydrocarboncracking reaction; and 2) by facilitating the removal of crackedequilibrium products (the hydrocarbon radicals) from the equilibrium (byincreasing the rate of hydrogenation), driving the equilibrium forwardin accordance with Le Chatelier's principle; 3) providing anotherreaction path to form hydrogenated cracked hydrocarbons; and 4)permitting the use of higher reaction temperatures without theconcomitant production of coke. The hydrogen sulfide directs theselectivity of the process away from producing coke by providinghydrogen atoms at an increased rate to the cracked hydrocarbon radicalsand by providing a thiol to the cracked hydrocarbon radicals—therebyinhibiting the cracked hydrocarbon radicals from agglomerating withother hydrocarbons.

As a result, the overall rate of the process of the present inventionmay be very high relative to other processes for cracking heavyhydrocarbon-containing feedstocks. The rate of the process may be atleast 350, or at least 400, or at least 500, or at least 600, or atleast 700, or at least 800, or at least 1000 kg/hr per m³ of reactionvolume up to 5000 kg/hr m³ of reaction volume.

Certain terms that are used herein are defined as follows: “Acridiniccompound” refers to a hydrocarbon compound including the structure:

As used in the present application, an acridinic compound includes anyhydrocarbon compound containing the above structure, including,naphthenic acridines, napththenic benzoacridines, and benzoacridines, inaddition to acridine.“Anaerobic conditions” means “conditions in which less than 0.5 vol. %oxygen as a gas is present”. For example, a process that occurs underanaerobic conditions, as used herein, is a process that occurs in thepresence of less than 0.5 vol. % oxygen in a gaseous form. Anaerobicconditions may be such that no detectable oxygen gas is present.“Aqueous” as used herein is defined as containing more than 50 vol. %water. For example, an aqueous solution or aqueous mixture, as usedherein, contains more than 50 vol. % water.“ASTM” refers to American Standard Testing and Materials.“Atomic hydrogen percentage” and “atomic carbon percentage” of ahydrocarbon-containing material—including crude oils, crude productssuch as syncrudes, bitumen, tar sands hydrocarbons, shale oil, crude oilatmospheric residues, crude oil vacuum residues, naphtha, kerosene,diesel, VGO, and hydrocarbons derived from liquefying coal—are asdetermined by ASTM Method D5291.“API Gravity” refers to API Gravity at 15.5° C., and as determined byASTM Method D6822.“Benzothiophenic compound” refers to a hydrocarbon compound includingthe structure:

As used in the present application, a benzothiophenic compound includesany hydrocarbon compound containing the above structure, includingdi-benzothiophenes, naphthenic-benzothiophenes,napththenic-di-benzothiophenes, benzo-naphtho-thiophenes,naphthenic-benzo-naphthothiophenes, and dinaphtho-thiophenes, inaddition to benzothiophene.“BET surface area” refers to a surface area of a material as determinedby ASTM Method D3663.“Blending” as used herein is defined to mean contact of two or moresubstances by intimately admixing the two or more substances.Boiling range distributions for a hydrocarbon-containing material may beas determined by ASTM Method D5307.“Bond” as used herein with reference to atoms in a molecule may refer toa covalent bond, a dative bond, or an ionic bond, dependent on thecontext.“Carbazolic compound” refers to a hydrocarbon compound including thestructure:

As used in the present application, a carbazolic compound includes anyhydrocarbon compound containing the above structure, includingnaphthenic carbazoles, benzocarbazoles, and napthenic benzocarbazoles,in addition to carbazole.“Carbon number” refers to the total number of carbon atoms in amolecule.“Catalyst” refers to a substance that increases the rate of a chemicalprocess and/or that modifies the selectivity of a chemical process asbetween potential products of the chemical process, where the substanceis not consumed by the process. A catalyst, as used herein, may increasethe rate of a chemical process by reducing the activation energyrequired to effect the chemical process. Alternatively, a catalyst, asused herein, may increase the rate of a chemical process by modifyingthe selectivity of the process between potential products of thechemical process, which may increase the rate of the chemical process byaffecting the equilibrium balance of the process. Further, a catalyst,as used herein, may not increase the rate of reactivity of a chemicalprocess but merely may modify the selectivity of the process as betweenpotential products.“Catalyst acidity by ammonia chemisorption” refers to the acidity of acatalyst substrate as measured by volume of ammonia adsorbed by thecatalyst substrate and subsequently desorbed from the catalyst substrateas determined by ammonia temperature programmed desorption between atemperature of 120° C. and 550° C. For clarity, a catalyst that isdecomposed in the measurement of acidity by ammonia temperatureprogrammed desorption to a temperature of 550° C. and/or a catalyst forwhich a measurement of acidity may not be determined by ammoniatemperature programmed desorption, e.g. a liquid or gas, is defined forpurposes of the present invention to have an indefinite acidity asmeasured by ammonia chemisorption. Ammonia temperature programmeddesorption measurement of the acidity of a catalyst is effected byplacing a catalyst sample that has not been exposed to oxygen ormoisture in a sample container such as a quartz cell; transferring thesample container containing the sample to a temperature programmeddesorption analyzer such as a Micrometrics TPD/TPR 2900 analyzer; in theanalyzer, raising the temperature of the sample in helium to 550° C. ata rate of 10° C. per minute; cooling the sample in helium to 120° C.;alternately flushing the sample with ammonia for 10 minutes and withhelium for 25 minutes a total of 3 times, and subsequently measuring theamount of ammonia desorbed from the sample in the temperature range from120° C. to 550° C. while raising the temperature at a rate of 10° C. perminute.“Coke” is a solid carbonaceous material that is formed primarily of ahydrocarbonaceous material and that is insoluble in toluene asdetermined by ASTM Method D4072.“Cracking” as used herein with reference to a hydrocarbon-containingmaterial refers to breaking hydrocarbon molecules in thehydrocarbon-containing material into hydrocarbon fragments, where thehydrocarbon fragments have a lower molecular weight than the hydrocarbonmolecule from which they are derived. Cracking conducted in the presenceof a hydrogen donor may be referred to as hydrocracking. Crackingeffected by temperature in the absence of a catalyst may be referred toa thermal cracking. Cracking may also produce some of the effects ofhydrotreating such as sulfur reduction, metal reduction, nitrogenreduction, and reduction of TAN.“Diesel” refers to hydrocarbons with a boiling range distribution from260° C. up to 343° C. (500° F. up to 650° F.) as determined inaccordance with ASTM Method D5307. Diesel content may be determined bythe quantity of hydrocarbons having a boiling range of from 260° C. to343° C. relative to a total quantity of hydrocarbons as measured byboiling range distribution in accordance with ASTM Method D5307.“Dispersible” as used herein with respect to mixing a solid, such as asalt, in a liquid is defined to mean that the components that form thesolid, upon being mixed with the liquid, are retained in the liquid atSTP for a period of at least 24 hours upon cessation of mixing the solidwith the liquid. A solid material is dispersible in a liquid if thesolid or its components are soluble in the liquid. A solid material isalso dispersible in a liquid if the solid or its components form acolloidal dispersion or a suspension in the liquid.“Distillate” or “middle distillate” refers to hydrocarbons with aboiling range distribution from 204° C. up to 343° C. (400° F. up to650° F.) as determined by ASTM Method D5307. Distillate may includediesel and kerosene.“Hydrogen” as used herein refers to molecular hydrogen unless specifiedas atomic hydrogen.“Insoluble” as used herein refers to a substance a majority (at least 50wt. %) of which does not dissolve or disperse in a liquid after a periodof 24 hours upon being mixed with the liquid at a specified temperatureand pressure, where the undissolved portion of the substance can berecovered from the liquid by physical means. For example, a fineparticulate material dispersed in a liquid is insoluble in the liquid if50 wt. % or more of the material may be recovered from the liquid bycentrifugation and filtration.“IP” refers to the Institute of Petroleum, now the Energy Institute ofLondon, United Kingdom.“Iso-paraffins” refer to branched chain saturated hydrocarbons.“Kerosene” refers to hydrocarbons with a boiling range distribution from204° C. up to 260° C. (400° F. up to 500° F.) at a pressure of 0.101MPa. Kerosene content may be determined by the quantity of hydrocarbonshaving a boiling range of from 204° C. to 260° C. at a pressure of 0.101MPa relative to a total quantity of hydrocarbons as measured by boilingrange distribution in accordance with ASTM Method D5307.“Lewis base” refers to a compound and/or material with the ability todonate one or more electrons to another compound.“Ligand” as used herein is defined as a molecule, compound, atom, or ionattached to, or capable of attaching to, a metal ion in a coordinationcomplex.“Light hydrocarbons” refers to hydrocarbons having a carbon number in arange from 1 to 6.“Mixing” as used herein is defined as contacting two or more substancesby intermingling the two or more substances. Blending, as used herein,is a subclass of mixing, where blending requires intimately admixing orintimately intermingling the two or more substances, for example into ahomogenous dispersion.“Monomer” as used herein is defined as a molecular compound or portionof a molecular compound that may be reactively joined with itself oranother monomer in repeated linked units to form a polymer.“Naphtha” refers to hydrocarbon components with a boiling rangedistribution from 38° C. up to 204° C. (100° F. up to 400° F.) at apressure of 0.101 MPa. Naphtha content may be determined by the quantityof hydrocarbons having a boiling range of from 38° C. to 204° C.relative to a total quantity of hydrocarbons as measured by boilingrange distribution in accordance with ASTM Method D5307. Content ofhydrocarbon components, for example, paraffins, iso-paraffins, olefins,naphthenes and aromatics in naphtha are as determined by ASTM MethodD6730.“Non-condensable gas” refers to components and/or a mixture ofcomponents that are gases at STP.“n-Paraffins” refer to normal (straight chain) saturated hydrocarbons.“Olefins” refer to hydrocarbon compounds with non-aromatic carbon-carbondouble bonds. Types of olefins include, but are not limited to, cis,trans, internal, terminal, branched, and linear.When two or more elements are described as “operatively connected”, theelements are defined to be directly or indirectly connected to allowdirect or indirect fluid flow between the elements.“Periodic Table” refers to the Periodic Table as specified by theInternational Union of Pure and Applied Chemistry (IUPAC), November2003. As used herein, an element of the Periodic Table of Elements maybe referred to by its symbol in the Periodic Table. For example, Cu maybe used to refer to copper, Ag may be used to refer to silver, W may beused to refer to tungsten etc.“Polyaromatic compounds” refer to compounds that include three or morearomatic rings. Examples of polyaromatic compounds include, but are notlimited anthracene and phenanthrene.“Polymer” as used herein is defined as a compound comprised ofrepetitively linked monomers.“Pore size distribution” refers a distribution of pore size diameters ofa material as measured by ASTM Method D4641.“SCFB” refers to standard cubic feet of gas per barrel of crude feed.“STP” as used herein refers to Standard Temperature and Pressure, whichis 25° C. and 0.101 MPa.The term “soluble” as used herein refers to a substance a majority (atleast 50 wt. %) of which dissolves in a liquid upon being mixed with theliquid at a specified temperature and pressure. For example, a materialdispersed in a liquid is soluble in the liquid if less than 50 wt. % ofthe material may be recovered from the liquid by centrifugation andfiltration.“TAN” refers to a total acid number expressed as milligrams (“mg”) ofKOH per gram (“g”) of sample. TAN is as determined by ASTM Method D664.“VGO” refers to hydrocarbons with a boiling range distribution of from343° C. up to 538° C. (650° F. up to 1000° F.) at 0.101 MPa. VGO contentmay be determined by the quantity of hydrocarbons having a boiling rangeof from 343° C. to 538° C. at a pressure of 0.101 MPa relative to atotal quantity of hydrocarbons as measured by boiling range distributionin accordance with ASTM Method D5307.“wppm” as used herein refers to parts per million, by weight.

The present invention is directed to a process for cracking ahydrocarbon-containing feedstock. A hydrocarbon-containing feedstockcontaining at least 20 wt. % of hydrocarbons having a boiling point ofgreater than 538° C. as determined in accordance with ASTM Method D5307is selected and is provided continuously or intermittently to a mixingzone at a selected rate. At least one metal-containing catalyst is alsoprovided to the mixing zone. Hydrogen is continuously or intermittentlyprovided to the mixing zone and blended with the hydrocarbon-containingfeedstock and the catalyst(s) in the mixing zone at temperature of from375° C. to 500° C. and at a total pressure of from 6.9 MPa to 27.5 MPa A(1000 psig to 4000 psig) to produce a vapor comprised of hydrocarbonsthat are vaporizable at the temperature and pressure within the mixingzone and a hydrocarbon-depleted feed residuum comprising hydrocarboncompounds that are liquid at the temperature and pressure within themixing zone. At least a portion of the vapor is separated from themixing zone while retaining the hydrocarbon-depleted feed residuum inthe mixing zone. Apart from the mixing zone, at least a portion of thevapor separated from the mixing zone may be condensed to produce aliquid hydrocarbon-containing product. Alternatively, apart from themixing zone, at least a portion of the vapor separated from the mixingzone may be hydrotreated by contacting the vapor with a hydrotreatingcatalyst and hydrogen at a temperature of from 260° C. to 425° C. and atotal pressure of from 3.4 MPa to 27.5 MPa (500 psig-4000 psig) toreduce sulfur, nitrogen, and olefinic hydrocarbons in the vapor.

The rate at which the hydrocarbon-containing feedstock is provided tothe mixing zone is selected to be at least 350 kg/hr per m³ of themixture volume in the mixing zone, where the mixture volume is definedby the combined volume of the hydrocarbon-depleted feed residuum (ifany), the catalyst(s), and the hydrocarbon-containing feedstock (if any)in the mixing zone. The rate at which the hydrocarbon-containingfeedstock is provided to the mixing zone may be sufficient to maintainat least a minimum mixture volume in the mixing zone. The mixing zonemay be located in a reactor, where the reactor has a reactor volume. Thehydrocarbon-containing feedstock and the catalyst(s) initially providedto the mixing zone may define an initial mixture volume, where theinitial mixture volume may be from 5% to 97% of the reactor volume. Therate at which the hydrocarbon-containing feedstock is provided to themixing zone may be sufficient to maintain the mixture volume of thecatalyst(s), hydrocarbon-depleted feed residuum, andhydrocarbon-containing feedstock at a level of at least 10%, or at least25%, or at least 40%, or at least 50%, or within 70%, or within 50%, orfrom 10% to 1940%, or from 15% to 1500%, or from 20% to 1000%, or from25% to 500%, or from 30% to 250%, or from 40% to 200% of the initialmixture volume.

Hydrocarbon-Containing Feedstock

The hydrocarbon-containing feedstock contains heavy hydrocarbons thatare subject to being cracked in the process. The hydrocarbon-containingfeedstock, therefore, is selected to contain at least 20 wt. %hydrocarbons having a boiling point of greater than 538° C. asdetermined in accordance with ASTM D5307. The hydrocarbon-containingfeedstock may be selected to contain at least 25 wt. %, or at least 30wt. %, or at least 35 wt. %, or at least 40 wt. %, or at least 45 wt. %,or at least 50 wt. % hydrocarbons having a boiling point of greater than538° C. as determined in accordance with ASTM Method D5307. Thehydrocarbon-containing feedstock may be selected to contain at least 20wt. % residue, or at least 25 wt. % residue, or at least 30 wt. %residue, or at least 35 wt. % residue, or at least 40 wt. % residue, orat least 45 wt. % residue, or least 50 wt. % residue.

The hydrocarbon-containing feedstock may contain significant quantitiesof lighter hydrocarbons as well as the heavy hydrocarbons. Thehydrocarbon-containing feedstock may contain at least 30 wt. %, or atleast 35 wt. %, or at least 40 wt. %, or at least 45 wt. %, or at least50 wt. % of hydrocarbons having a boiling point of 538° C. or lessdetermined in accordance with ASTM Method D5307. Thehydrocarbon-containing feedstock may contain at least 20 wt. %, or atleast 25 wt. %, or at least 30 wt. %, or at least 35 wt. %, or at least40 wt. %, or at least 45 wt. % of naphtha and distillate hydrocarbons.The hydrocarbon-containing feedstock may be a crude oil, or may be atopped crude oil.

The hydrocarbon-containing feedstock may also contain quantities ofmetals such as vanadium and nickel. The hydrocarbon-containing feedstockmay contain at least 50 wppm vanadium and at least 20 wppm nickel.

The hydrocarbon-containing feedstock may also contain quantities ofsulfur and nitrogen. The hydrocarbon containing feedstock may contain atleast 2 wt. % sulfur, or at least 3 wt. % sulfur; and thehydrocarbon-containing feedstock may contain at least 0.25 wt. %nitrogen, or at least 0.4 wt. % nitrogen.

The hydrocarbon-containing feedstock may also contain appreciablequantities of naphthenic acids. For example, the hydrocarbon-containingfeedstock may have a TAN of at least 0.5, or at least 1.0, or at least2.0.

The process of the present invention is particularly applicable tocertain heavy petroleum and coal derived hydrocarbon-containingfeedstocks. The hydrocarbon-containing feedstock may be a heavy or anextra-heavy crude oil containing significant quantities of residue orpitch; a topped heavy or topped extra-heavy crude oil containingsignificant quantities of residue or pitch; bitumen; hydrocarbonsderived from tar sands; shale oil; crude oil atmospheric residues; crudeoil vacuum residues; asphalts; and hydrocarbons derived from liquefyingcoal.

In the process of the present invention, an initial charge ofhydrocarbon-containing feedstock is provided to a mixing zone for mixingand reaction with hydrogen and one or more catalyst(s). Additionalhydrocarbon-containing feedstock is then provided continuously orintermittently to the mixing zone at a rate of at least 350 kg/hr per m³of the mixture volume, where the mixture volume is as defined above.

Hydrogen

The hydrogen that is mixed with the hydrocarbon-containing feedstock andthe catalyst in the process of the present invention is derived from ahydrogen source. The hydrogen source may be hydrogen gas obtained fromany conventional sources or methods for producing hydrogen gas.Optionally, the hydrogen may provided in a synthesis gas.

Catalyst

One or more metal-containing catalysts may be utilized in the process ofthe present invention. The one or more metal-containing catalysts areselected to catalyze hydrocracking of the hydrocarbon-containingfeedstock. Each catalyst utilized in the process of the presentinvention preferably has little or no acidity to avoid catalyzing theformation of hydrocarbon radical cations and thereby avoid catalyzingthe formation of coke. Each catalyst utilized in the process of theinvention preferably has an acidity as measured by ammonia chemisorptionof at most 200, or at most 100, or at most 50, or at most 25, or at most10 μmol ammonia per g of catalyst, and most preferably has an acidity asmeasured by ammonia chemisorption of 0 μmol ammonia per g of catalyst.In an embodiment, the one or more catalysts comprise at most 0.1 wt. %,or at most 0.01 wt. %, or at most 0.001 wt. % of alumina,alumina-silica, or silica, and, preferably, the one or more catalystscontain no detectable alumina, alumina-silica, or silica.

The one or more metal-containing catalysts used in the process of thepresent invention may contain little or no oxygen. The catalyticactivity of the metal-containing catalyst(s) in the process of thepresent invention is, in part, believed to be due to the availability ofelectrons from the catalyst(s) to stabilize cracked molecules in thecrude oil. Due to its electronegativity, oxygen tends to reduce theavailability of electrons from a catalyst when it is present in thecatalyst in appreciable quantities, therefore, each catalyst utilized inthe process preferably contains little or no oxygen. Each catalystutilized in the process may comprise at most 0.1 wt. %, or at most 0.05wt. %, or at most 0.01 wt. % oxygen as measured by neutron activation.

One or more of the metal-containing catalysts may be a solid particulatesubstance having a particle size distribution with a relatively smallmean and/or median particle size, where the solid catalyst particlespreferably are nanometer size particles. A catalyst may have a particlesize distribution with a median particle size and/or mean particle sizeof at least 50 nm, or at least 75 nm, or up to 5 nm, or up to 1 μm; orup to 750 nm, or from 50 nm up to 5 μm. A solid particulate catalysthaving a particle size distribution with a large quantity of smallparticles, for example having a mean and/or median particle size of upto 5 μm, has a large aggregate surface area since little of thecatalytically active components of the catalyst are located within theinterior of a particle. A particulate catalyst having a particle sizedistribution with a large quantity of small particles, therefore, may bedesirable for use in the process of the present invention to provide arelatively high degree catalytic activity due to the surface area of thecatalyst available for catalytic activity. A catalyst used in theprocess of the invention may be a solid particulate substance preferablyhaving a particle size distribution with a mean particle size and/ormedian particle size of up to 1 μm, preferably having a pore sizedistribution with a mean pore diameter and/or a median pore diameter offrom 50 to 1000 angstroms, or from 60 to 350 angstroms, preferablyhaving a pore volume of at least 0.2 cm³/g, or at least 0.25 cm³/g or atleast 0.3 cm³/g, or at least 0.35 cm³/g, or at least 0.4 cm³/g, andpreferably having a BET surface area of at least 50 m²/g, or at least100 m²/g, and up to 400 m²/g, or up to 500 m²/g.

A solid particulate catalyst utilized in the process of the presentinvention may be insoluble in the hydrocarbon-containing feed and in thehydrocarbon-depleted feed residuum formed by the process of the presentinvention. A solid particulate catalyst having a particle sizedistribution with a median and/or mean particle size of at least 50 nmmay be insoluble in the hydrocarbon-containing feed and thehydrocarbon-depleted residuum due, in part, to the size of theparticles, which may be too large to be solvated by thehydrocarbon-containing feed or the residuum Use of a solid particulatecatalyst which is insoluble in the hydrocarbon-containing feed and thehydrocarbon-depleted feed residuum may be desirable in the process ofthe present invention so that the catalyst may be separated from theresiduum formed by the process, and subsequently regenerated for reusein the process.

A catalyst that may be used in the process of the present invention hasan acidity as measured by ammonia chemisorption of at most 200 μmolammonia per gram of catalyst, and comprises a material comprised of ametal of Column(s) 6-10 of the Periodic Table or a compound of a metalof Column(s) 6-10 of the Periodic Table. The catalyst may be abi-metallic catalyst comprised of a metal of Column 6, 14, or 15 of thePeriodic Table or a compound of a metal of Column 6, 14, or 15 of thePeriodic Table and a metal of Column(s) 3 or 7-15 of the Periodic Tableor a compound of a metal of Column(s) 3 or 7-15 of the Periodic Table,where the catalyst has an acidity as measured by ammonia chemisorptionof at most 200 μmol ammonia per g of catalyst.

In a preferred embodiment, a catalyst that is mixed with thehydrocarbon-containing feedstock and the hydrogen in the mixing zone iscomprised of a material that is comprised of a first metal, a secondmetal, and sulfur. The first metal of the material of the catalyst maybe a metal selected from the group consisting of copper (Cu), iron (Fe),bismuth (Bi), nickel (Ni), cobalt (Co), silver (Ag), manganese (Mn),zinc (Zn), tin (Sn), ruthenium (Ru), lanthanum (La), cerium (Ce),praseodymium (Pr), samarium (Sm), europium (Eu), ytterbium (Yb),lutetium (Lu), dysprosium (Dy), lead (Pb), and antimony (Sb). The firstmetal may be relatively electron-rich, inexpensive, and relativelynon-toxic, and preferably the first metal is selected to be copper oriron, most preferably copper. The second metal of the material of thecatalyst is a metal selected from the group consisting of molybdenum(Mo), tungsten (W), vanadium (V), tin (Sn), and antimony (Sb), where thesecond metal is not the same metal as the first metal.

The material of a preferred catalyst may be comprised of at least threelinked chain elements, where the chain elements are comprised of a firstchain element and a second chain element. The first chain elementincludes the first metal and sulfur and has a structure according toformula (I) and the second chain element includes the second metal andsulfur and has a structure according to formula (II):

where M¹ is the first metal and M² is the second metal. The catalystmaterial containing the chain elements contains at least one first chainelement and at least one second chain element. The chain elements of thematerial of the catalyst are linked by bonds between the two sulfuratoms of a chain element and the metal of an adjacent chain element. Achain element of the material of the catalyst may be linked to one, ortwo, or three, or four other chain elements, where each chain elementmay be linked to other chain elements by bonds between the two sulfuratoms of a chain element and the metal of an adjacent chain element. Atleast three linked chain elements may be sequentially linked in series.At least a portion of the material of the catalyst containing the chainelements may be comprised of the first metal and the second metal linkedby, and bonded to, sulfur atoms according to formula (III):

where M¹ is the first metal, M² is the second metal, and x is at least2. The material of the catalyst may be a polythiometallate polymer,where each monomer of the polymer is the structure as shown in formula(III) where x=1, and the polythiometallate polymer is the structure asshown in formula (III) where x is at least 5. At least a portion of thematerial of the catalyst may be comprised of the first metal and secondmetal, where the first metal is linked to the second metal by sulfuratoms as according to formula (IV) or formula (V):

where M¹ is the first metal and where M² is the second metal.

The material of the catalyst described above may comprise a third chainelement comprised of sulfur and a third metal selected from the groupconsisting of Cu, Fe, Bi, Ag, Mn, Zn, Ni, Co, Sn, Re, Rh, Pd, Ir, Pt,Ce, La, Pr, Sm, Eu, Yb, Lu, Dy, Pb, Cd, Sb, and In, where the thirdmetal is not the same as the first metal or the second metal. The thirdchain element has a structure according to formula (VI):

where M³ is the third metal. If the material of the catalyst contains athird chain element, at least a portion of the third chain element ofthe material of the catalyst is linked by bonds between the two sulfuratoms of a chain element and the metal of an adjacent chain element.

At least a portion of the catalyst material may be comprised of thefirst metal, the second metal, and sulfur having a structure accordingto formula (VII):

where M is either the first metal or the second metal, and at least oneM is the first metal and at least one M is the second metal. Thecatalyst material as shown in formula (VII) may include a third metalselected from the group consisting of Cu, Fe, Bi, Ag, Mn, Zn, Ni, Co,Sn, Re, Rh, Pd, Ir, Pt, Ce, La, Pr, Sm, Eu, Yb, Lu, Dy, Pb, Cd, Sb, andIn, where the third metal is not the same as the first metal or thesecond metal, and where M is either the first metal, or the secondmetal, or the third metal, and at least one M is the first metal, atleast one M is the second metal, and at least one M is the third metal.The portion of the catalyst material comprised of the first metal, thesecond metal, and sulfur may also have a structure according to formula(VIII):

where M is either the first metal or the second metal, at least one M isthe first metal and at least one M is the second metal, and x is atleast 2.

At least a portion of the material of the catalyst may be comprised ofthe first metal, the second metal, and sulfur having a structureaccording to formula (IX):

where M is either the first metal or the second metal, at least one M isthe first metal and at least one M is the second metal, and X isselected from the group consisting of SO₄, PO₄, oxalate (C₂O₄),acetylacetonate, acetate, citrate, tartrate, Cl, Br, I, ClO₄, and NO₃.For example, the material of the catalyst may contain copperthiometallate-sulfate having the structure shown in formula (X):

where n may be an integer greater than or equal to 1. The material ofthe catalyst as shown in formula (IX) may include a third metal selectedfrom the group consisting of Cu, Fe, Bi, Ag, Mn, Zn, Ni, Co, Sn, Re, Rh,Pd, Ir, Pt, Ce, La, Pr, Sm, Eu, Yb, Lu, Dy, Pb, Cd, Sb, and In, wherethe third metal is not the same as the first metal or the second metal,where M is either the first metal, or the second metal, or the thirdmetal, and at least one M is the first metal, at least one M is thesecond metal, and at least one M is the third metal. The portion of thematerial of the catalyst comprised of the first metal, the second metal,and sulfur may also have a polymeric structure according to formula(XI):

where M is either the first metal or the second metal, at least one M isthe first metal and at least one M is the second metal, X is selectedfrom the group consisting of SO₄, PO₄, oxalate (C₂O₄), acetylacetonate,acetate, citrate, tartrate, Cl, Br, I, ClO₄, and NO₃, and x is at least2 and preferably is at least 5;

At least a portion of the catalyst material may be comprised of thefirst metal, the second metal, and sulfur having a structure accordingto formula (XII):

where M is either the first metal or the second metal, at least one M isthe first metal and at least one M is the second metal, and X isselected from the group consisting of SO₄, PO₄, oxalate (C₂O₄),acetylacetonate, acetate, citrate, tartrate, Cl, Br, I, ClO₄, and NO₃.The material of the catalyst as shown in formula (XII) may include athird metal selected from the group consisting of Cu, Fe, Bi, Ag, Mn,Zn, Ni, Co, Sn, Re, Rh, Pd, Ir, Pt, Ce, La, Pr, Sm, Eu, Yb, Lu, Dy, Pb,Cd, Sb, and In, where the third metal is not the same as the first metalor the second metal, and where M is either the first metal, or thesecond metal, or the third metal, and at least one M is the first metal,at least one M is the second metal, and at least one M is the thirdmetal. The portion of the catalyst material comprised of the firstmetal, the second metal, and sulfur may also have a polymeric structureaccording to formula (XIII).

where M is either the first metal or the second metal, and at least oneM is the first metal and at least one M is the second metal, X isselected from the group consisting of SO₄, PO₄, oxalate (C₂O₄),acetylacetonate, acetate, citrate, tartrate, Cl, Br, I, ClO₄, and NO₃,and x is at least 2 and preferably is at least 5.

At least a portion of the catalyst material may be comprised of thefirst metal, the second metal, and sulfur having a structure accordingto formula (XIV):

where M is either the first metal or the second metal, at least one M isthe first metal and at least one M is the second metal, and X isselected from the group consisting of SO₄, PO₄, oxalate (C₂O₄),acetylacetonate, acetate, citrate, tartrate, Cl, Br, I, ClO₄, and NO₃.For example, at least a portion of the catalyst material may have astructure in accordance with formula (XV):

where X is selected from the group consisting of SO₄, PO₄, oxalate(C₂O₄), acetylacetonate, acetate, citrate, tartrate, Cl, Br, I, ClO₄,and NO₃, and n is an integer equal to or greater than 1. The catalystmaterial as shown in formula (XIV) may include a third metal selectedfrom the group consisting of Cu, Fe, Bi, Ag, Mn, Zn, Ni, Co, Sn, Re, Rh,Pd, Ir, Pt, Ce, La, Pr, Sm, Eu, Yb, Lu, Dy, Pb, Cd, Sb, and In, wherethe third metal is not the same as the first metal or the second metal,and where M is either the first metal, or the second metal, or the thirdmetal, and at least one M is the first metal, at least one M is thesecond metal, and at least one M is the third metal. The portion of thecatalyst material comprised of the first metal, the second metal, andsulfur may also have a polymeric structure according to formula (XVI):

where M is either the first metal or the second metal, at least one M isthe first metal and at least one M is the second metal, X is selectedfrom the group consisting of SO₄, PO₄, oxalate (C₂O₄), acetylacetonate,acetate, citrate, tartrate, Cl, Br, I, ClO₄, and NO₃, and x is at least2 and preferably is at least 5.

A preferred catalyst preferably is formed primarily of a materialcomprised of the first metal, second metal, and sulfur as describedabove, and the material of the preferred catalyst is formed primarily ofthe first metal, second metal, and sulfur as described above. The firstmetal, second metal, and sulfur may comprise at least 75 wt. %, or atleast 80 wt. %, or at least 85 wt. %, or at least 90 wt. %, or at least95 wt. %, or at least 99 wt. % or 100 wt. % of the material of thecatalyst structured as described above, where the material of thecatalyst comprises at least 50 wt. % or at least 60 wt. %, or at least70 wt. %, or at least 75 wt. %, or at least 80 wt. %, or at least 90 wt.%, or at least 95 wt. %, or at least 99 wt. % or 100 wt. % of thecatalyst.

The first metal may be present in the material of a preferred catalystdescribed above, in an atomic ratio relative to the second metal of atleast 1:2. The atomic ratio of the first metal to the second metal inthe material of the catalyst, and/or in the catalyst, may be greaterthan 1:2, or at least 2:3, or at least 1:1, or at least 2:1, or at least3:1, or at least 5:1. It is believed that the first metal contributessignificantly to the catalytic activity of the catalyst in the processof the present invention when the first metal is present in the materialof the catalyst, and/or in the catalyst, in an amount relative to thesecond metal ranging from slightly less of the first metal to the secondmetal to significantly more of the first metal to the second metal.Therefore, the first metal may be incorporated in the material of thecatalyst, and/or in the catalyst, in an amount, relative to the secondmetal, such that the atomic ratio of the first metal to the second metalranges from one half to significantly greater than one, such that thefirst metal is not merely a promoter of the second metal in thecatalyst.

A preferred catalyst—when primarily formed of the material of thecatalyst, where the material of the catalyst is primarily formed of thefirst metal, the second metal, and sulfur structured as described above,and particularly when the first metal, the second metal, and the sulfurthat form the material of the catalyst are not supported on a carrier orsupport material to form the catalyst—may have a significant degree ofporosity, pore volume, and surface area. In the absence of a support ora carrier, the catalyst may have a pore size distribution, where thepore size distribution has a mean pore diameter and/or a median porediameter of from 50 angstroms to 1000 angstroms, or from 60 angstroms to350 angstroms. In the absence of a support or a carrier, the catalystmay have a pore volume of at least 0.2 cm³/g, or at least 0.25 cm³/g, orat least 0.3 cm³/g, or at least 0.35 cm³/g, or at least 0.4 cm³/g. Inthe absence of a support or a carrier, the catalyst may have a BETsurface area of at least 50 m²/g, or at least 100 m², and up to 400 m²/gor up to 500 m²/g.

The relatively large surface area of the preferred catalyst,particularly relative to conventional non-supported bulk metalcatalysts, is believed to be due, in part, to the porosity of thecatalyst imparted by at least a portion of the material of the catalystbeing formed of abutting or adjoining linked tetrahedrally structuredatomic formations of the first metal and sulfur and the second metal andsulfur, where the tetrahedrally structured atomic formations may beedge-bonded. Interstices or holes that form the pore structure of thecatalyst may be present in the material of the catalyst as a result ofthe bonding patterns of the tetrahedral structures. Preferred catalysts,therefore, may be highly catalytically active since 1) the catalystshave a relatively large surface area; and 2) the surface area of thecatalysts is formed substantially, or entirely, of the elements thatprovide catalytic activity—the first metal, the second metal, andsulfur.

The material of a preferred catalyst may contain less than 0.5 wt. % ofligands other than sulfur-containing ligands. Ligands, other thansulfur-containing ligands, may not be present in significant quantitiesin the material since they may limit the particle size of the materialof the catalyst to less than 50 nm, for example, by inhibiting the firstmetal and the second metal from forming sulfur-bridged chains.

Method of Preparing Preferred Catalysts

A preferred catalyst utilized in the process of the present inventionmay be prepared by mixing a first salt and a second salt in an aqueousmixture under anaerobic conditions at a temperature of from 15° C. to150° C., and separating a solid from the aqueous mixture to produce thecatalyst material.

The first salt utilized to form a preferred catalyst includes a cationiccomponent comprising a metal in any non-zero oxidation state selectedfrom the group consisting of Cu, Fe, Ni, Co, Bi, Ag, Mn, Zn, Sn, Ru, La,Ce, Pr, Sm, Eu, Yb, Lu, Dy, Pb, and Sb, where the metal of the cationiccomponent is the first metal of the material of the catalyst. Thecationic component of the first salt may consist essentially of a metalselected from the group consisting of Cu, Fe, Bi, Ni, Co, Ag, Mn, Zn,Sn, Ru, La, Ce, Pr, Sm, Eu, Yb, Lu, Dy, Pb, and Sb. The cationiccomponent of the first salt must be capable of bonding with the anioniccomponent of the second salt to form the material of the catalyst in theaqueous mixture at a temperature of from 15° C. to 150° C. and underanaerobic conditions.

The first salt also contains an anionic component associated with thecationic component of the first salt to form the first salt. The anioniccomponent of the first salt may be selected from a wide range ofcounterions to the cationic component of the first salt so long as thecombined cationic component and the anionic component of the first saltform a salt that is dispersible, and preferably soluble, in the aqueousmixture in which the first salt and the second salt are mixed, and solong as the anionic component of the first salt does not prevent thecombination of the cationic component of the first salt with the anioniccomponent of the second salt in the aqueous mixture to form the materialof the catalyst. The anionic component of the first salt may be selectedfrom the group consisting of sulfate, chloride, bromide, iodide,acetate, acetylacetonate, phosphate, nitrate, perchlorate, oxalate,citrate, and tartrate.

The anionic component of the first salt may associate with or beincorporated into a polymeric structure including the cationic componentof the first salt and the anionic component of the second salt to formthe material of the catalyst. For example, the anionic component of thefirst salt may complex with a polymeric structure formed of the cationiccomponent of the first salt and the anionic component of the second saltas shown in formulas (XI) and (XIII) above, where X=the anioniccomponent of the first salt, or may be incorporated into a polymericstructure including the cationic component of the first salt and theanionic component of the second salt as shown in formula (XVI) above,where X=the anionic component of the first salt.

Certain compounds are preferred for use as the first salt to form apreferred catalyst. In particular, the first salt is preferably selectedfrom the group consisting of CuSO₄, copper acetate, copperacetylacetonate, FeSO₄, Fe₂(SO₄)₃, iron acetate, iron acetylacetonate,NiSO₄, nickel acetate, nickel acetylacetonate, CoSO₄, cobalt acetate,cobalt acetylacetonate, ZnCl₂, ZnSO₄, zinc acetate, zincacetylacetonate, silver acetate, silver acetylacetonate, SnSO₄, SnCl₄,tin acetate, tin acetylacetonate, MnSO₄, manganese acetate, manganeseacetylacetonate, bismuth acetate, bismuth acetylacetonate, and hydratesthereof. These materials are generally commercially available, or may beprepared from commercially available materials according to well-knownmethods.

The first salt is contained in an aqueous solution or an aqueousmixture, where the aqueous solution or aqueous mixture containing thefirst salt (hereinafter the “first aqueous solution”) is mixed with anaqueous solution or an aqueous mixture containing the second salt(hereinafter the “second aqueous solution”) in the aqueous mixture toform the material of the preferred catalyst. The first salt may bedispersible, and most preferably soluble, in the first aqueous solutionand is dispersible, and preferably soluble, in the aqueous mixture ofthe first and second salts. The first aqueous solution may contain morethan 50 vol. % water, or at least 75 vol. % water, or at least 90 vol. %water, or at least 95 vol. % water, and may contain more than 0 vol. %but less than 50 vol. %, or at most 25 vol. %, or at most 10 vol. %, orat most 5 vol. % of an organic solvent containing from 1 to 5 carbonsselected from the group consisting of an alcohol, a diol, an aldehyde, aketone, an amine, an amide, a furan, an ether, acetonitrile, andmixtures thereof. The organic solvent present in the first aqueoussolution, if any, should be selected so that the organic compounds inthe organic solvent do not inhibit reaction of the cationic component ofthe first salt with the anionic component of the second salt uponforming an aqueous mixture containing the first and second salts, e.g.,by forming ligands or by reacting with the first or second salts ortheir respective cationic or anionic components. The first aqueoussolution may contain no organic solvent, and may consist essentially ofwater, preferably deionized water, and the first salt.

The concentration of the first salt in the first aqueous solution may beselected to promote formation of a preferred catalyst having a particlesize distribution with a small mean and/or median particle size, wherethe particles have a relatively large surface area, upon mixing thefirst salt and the second salt in the aqueous mixture. To promote theformation of a catalyst material having a relatively large surface areaand having particle size distribution with a relatively small meanand/or median particle size, the first aqueous solution may contain atmost 3 moles per liter, or at most 2 moles per liter, or at most 1 moleper liter, or at most 0.6 moles per liter, or at most 0.2 moles perliter of the first salt.

The second salt utilized to form a preferred catalyst includes ananionic component that is a tetrathiometallate of molybdenum, vanadium,tungsten, tin or antimony. In particular, the second salt may contain ananionic component that is selected from the group consisting of MoS₄ ²⁻,WS₄ ²⁻, VS₄ ³⁻, SnS₄ ⁴⁻, and SbS₄ ³⁻.

The second salt also contains a cationic component associated with theanionic component of the second salt to form the second salt. Thecationic component of the second salt may be selected from an ammoniumcounterion, and alkali metal and alkaline earth metal counterions to thetetrathiometallate anionic component of the second salt so long as thecombined cationic component and the anionic component of the second saltform a salt that is dispersable, and preferably soluble, in the aqueousmixture in which the first salt and the second salt are mixed, and solong as the cationic component of the second salt does not prevent thecombination of the cationic component of the first salt with the anioniccomponent of the second salt in the aqueous mixture to form the catalystmaterial. The cationic component of the second salt may comprise one ormore sodium ions, or one or more potassium ions, or one or more ammoniumions.

Certain compounds are preferred for use as the second salt used to formthe material of the catalyst and/or the catalyst. In particular, thesecond salt is preferably selected from the group consisting of Na₂MoS₄,Na₂WS₄, Na₃ VS₄, K₂MoS₄, K₂WS₄, K₃VS₄, (NH₄)₂MoS₄, (NH₄)₂WS₄, (NH₄)₃VS₄,Na₄SnS₄, (NH₄)₄SnS₄, (NH₄)₃SbS₄, Na₃SbS₄, and hydrates thereof.

The second salt may be a commercially available tetrathiomolybdate ortetrathiotungstate salt. For example, the second salt may be ammoniumtetrathiomolybdate, which is commercially available from AAA MolybdenumProducts, Inc. 7233 W. 116 Pl., Broomfield, Colo., USA 80020, ammoniumtetrathiotungstate, which is commercially available from Sigma-Aldrich,3050 Spruce St., St. Louis, Mo., USA 63103, or ammoniumtetrathiovanadate, which is commercially available from Chemos GmbH,Germany

Alternatively, the second salt may be produced from a commerciallyavailable tetrathiomolybdate or tetrathiotungstate salt. For example,the second salt may be produced from a ammonium tetrathiomolybdate,ammonium tetrathiotungstate, or from ammonium tetrathiovanadate salt.The second salt may be formed from the commercially available ammoniumtetrathiometallate salts by exchanging the cationic ammonium componentof the commercially available salt with a desired alkali or alkalineearth cationic component from a separate salt. The exchange of thecationic components to form the desired second salt may be effected bymixing the commercially available salt and the salt containing thedesired cationic component in an aqueous solution to form the desiredsecond salt.

A method of forming the second salt is to disperse an ammoniumtetrathiomolybdate, ammonium tetrathiotungstate, or ammoniumtetrathiovanadate in an aqueous solution, preferably water, and todisperse an alkali metal or alkaline earth metal cationic componentdonor salt, preferably a carbonate, in the aqueous solution, where thecationic component donor salt is provided in an amount relative to theammonium tetrathiomolybdate, ammonium tetrathiotungstate, or ammoniumtetrathiovanadate salt to provide a stoichiometrially equivalent orgreater amount of its cation to ammonium of the ammoniumtetrathiomolybdate, ammonium tetrathiotungstate, or ammoniumtetrathiovanadate salt. The aqueous solution may be heated to atemperature of at least 50° C., or at least 65° C. up to 100° C. toevolve ammonia from the ammonium containing salt and carbon dioxide fromthe carbonate containing salt as gases, and to form the second salt. Forexample a Na₂MoS₄ salt may be prepared for use as the second salt bymixing commercially available (NH₄)₂MoS₄ and Na₂CO₃ in water at atemperature of 70° C.-80° C. for a time period sufficient to permitevolution of a significant amount, preferably substantially all, ofammonia and carbon dioxide gases from the solution, typically from 30minutes to 4 hours, and usually about 2 hours.

If the second salt is a sodium tetrathiostannate salt, it may beproduced by dissolving Na₂Sn(OH)₆ and Na₂S in a 1:4 molar ratio inboiling deionized water (100 g of Na₂Sn(OH)₆ per 700 ml of water and 250g of Na₂S per 700 ml of water), stirring the mixture at 90-100° C. for2-3 hours, adding finely pulverized MgO to the mixture at a 2:5 wt.ratio relative to the Na₂Sn(OH)₆ and continuing stirring the mixture at90-100° C. for an additional 2-3 hours, cooling and collectingprecipitated impurities from the mixture, then concentrating theremaining solution by 50-60 vol. %, allowing the concentrated solutionto stand, then collecting the Na₄SnS₄ that crystallizes from theconcentrated solution. An ammonium tetrathiostannate salt may beproduced by mixing SnS₂ with (NH₄)₂S in a 1:2 mole ratio in liquidammonia under an inert gas (e.g. nitrogen), filtering, and recoveringthe solid (NH)₄SnS₄ as a residue.

The second salt is contained in an aqueous solution (the second aqueoussolution, as noted above), where the second aqueous solution containingthe second salt is mixed with the first aqueous solution containing thefirst salt in the aqueous mixture to form the preferred catalyst. Thesecond salt is preferably dispersible, and most preferably soluble, inthe second aqueous solution and is dispersible, and preferably soluble,in the aqueous mixture containing the first and second salts. The secondaqueous solution contains more than 50 vol. % water, or at least 75 vol.% water, or at least 90 vol. % water, or at least 95 vol. % water, andmay contain more than 0 vol. % but less than 50 vol. %, or at most 25vol. %, or at most 10 vol. %, or at most 5 vol. % of an organic solventcontaining from 1 to 5 carbons and selected from the group consisting ofan alcohol, a diol, an aldehyde, a ketone, an amine, an amide, a furan,an ether, acetonitrile, and mixtures thereof. The organic solventpresent in the second aqueous solution, if any, should be selected sothat the organic compounds in the organic solvent do not inhibitreaction of the cationic component of the first salt with the anioniccomponent of the second salt upon forming an aqueous mixture containingthe first and second salts, e.g., by forming ligands or by reacting withthe first or second salts or their respective cationic or anioniccomponents. Preferably, the second aqueous solution contains no organicsolvent. Most preferably the second aqueous solution consistsessentially of water, preferably deionized, and the second salt.

The concentration of the second salt in the second aqueous solution maybe selected to promote formation of a catalyst having a particle sizedistribution with a small mean and/or median particle size and having arelatively large surface area per particle upon mixing the first saltand the second salt in the aqueous mixture. To promote the formation ofa catalyst material having a particle size distribution with arelatively small mean and/or median particle size, the second aqueoussolution may contain at most 0.8 moles per liter, or at most 0.6 molesper liter, or at most 0.4 moles per liter, or at most 0.2 moles perliter, or at most 0.1 moles per liter of the second salt.

The first and second solutions containing the first and second salts,respectively, are mixed in an aqueous mixture to form the preferredcatalyst. The amount of the first salt relative to the amount of thesecond salt provided to the aqueous mixture may be selected so that theatomic ratio of the cationic component metal of the first salt to themetal of the anionic component of the second salt is at least 1:2, orgreater than 1:2, or at least 2:3, or at least 1:1, and at most 20:1, orat most 15:1, or at most 10:1.

The aqueous mixture of the first and second salts is formed by addingthe first aqueous solution containing the first salt and the secondaqueous solution containing the second salt into an aqueous solutionseparate from both the first aqueous solution and the second aqueoussolution. The separate aqueous solution will be referred hereafter asthe “third aqueous solution”. The third aqueous solution may containmore than 50 vol. % water, or at least 75 vol. % water, or at least 90vol. % water, or at least 95 vol. % water, and may contain more than 0vol. % but less than 50 vol. %, or at most 25 vol. %, or at most 10 vol.%, or at most 5 vol. % of an organic solvent containing from 1 to 5carbons and selected from the group consisting of an alcohol, a diol, analdehyde, a ketone, an amine, an amide, a furan, an ether, acetonitrile,and mixtures thereof. The organic solvent present in the third aqueoussolution, if any, should be selected so that the organic compounds inthe organic solvent do not inhibit reaction of the cationic component ofthe first salt with the anionic component of the second salt uponforming the aqueous mixture, e.g., by forming ligands or reacting withthe cationic component of the first salt or with the anionic componentof the second salt. Preferably, the third aqueous solution contains noorganic solvent, and most preferably comprises deionized water.

The aqueous mixture of the first and second salts is formed by combiningthe first aqueous solution containing the first salt and the secondaqueous solution containing the second salt in the third aqueoussolution. The volume ratio of the third aqueous solution to the firstaqueous solution containing the first salt may be from 0.5:1 to 50:1where the first aqueous solution may contain at most 3, or at most 2, orat most 1, or at most 0.8, or at most 0.5, or at most 0.3 moles of thefirst salt per liter of the first aqueous solution. Likewise, the volumeratio of the third aqueous solution to the second aqueous solutioncontaining the second salt may be from 0.5:1 to 50:1 where the secondaqueous solution may contain at most 0.8, or at most 0.4, or at most0.2, or at most 0.1 moles of the second salt per liter of the secondaqueous solution.

The first salt and the second salt may be combined in the aqueousmixture so that the aqueous mixture containing the first and secondsalts contains at most 1.5, or at most 1.2, or at most 1, or at most0.8, or at most 0.6 moles of the combined first and second salts perliter of the aqueous mixture. The particle size of the catalyst materialproduced by mixing the first and second salts in the aqueous mixtureincreases, and the surface area of the particles decreases, withincreasing concentrations of the salts. Therefore, to limit the particlesizes in the particle size distribution of the catalyst material and toincrease the relative surface area of the particles, the aqueous mixturemay contain at most 0.8 moles of the combined first and second salts perliter of the aqueous mixture, more preferably at most 0.6 moles, or atmost 0.4 moles, or at most 0.2 moles of the combined first and secondsalts per liter of the aqueous mixture. The amount of the first salt andthe total volume of the aqueous mixture may be selected to provide atmost 1, or at most 0.8, or at most 0.4 moles of the cationic componentof the first salt per liter of the aqueous mixture and the amount of thesecond salt and the total volume of the aqueous mixture may be selectedto provide at most 0.4, or at most 0.2, or at most 0.1, or at most 0.01moles of the anionic component of the second salt per liter of theaqueous mixture.

The rate of addition of the first and second aqueous solutionscontaining the first and second salts, respectively, to the aqueousmixture may be controlled to limit the instantaneous concentration ofthe first and second salts in the aqueous mixture to produce a catalystmaterial comprised of relatively small particles having relatively largesurface area Limiting the instantaneous concentration of the salts inthe aqueous mixture may reduce the mean and/or median particle size ofthe resulting catalyst material by limiting the simultaneousavailability of large quantities of the cationic components of the firstsalt and large quantities of the anionic components of the second saltthat may interact to form a catalyst material comprised primarily ofrelatively large particles. The rate of addition of the first and secondsolutions to the aqueous mixture may be controlled to limit theinstantaneous concentration of the first salt and the second salt in theaqueous mixture to at most 0.05 moles per liter, or at most 0.01 molesper liter, or at most 0.001 moles per liter.

The first aqueous solution containing the first salt and the secondaqueous solution containing the second salt may be added to the thirdaqueous solution, preferably simultaneously, at a controlled rateselected to provide a desired instantaneous concentration of the firstsalt and the second salt in the aqueous mixture. The first aqueoussolution containing the first salt and the second aqueous solutioncontaining the second salt may be added to the third aqueous solution ata controlled rate by adding the first aqueous solution and the secondaqueous solution to the third aqueous solution in a dropwise manner Therate that drops of the first aqueous solution and the second aqueoussolution are added to the third aqueous solution may be controlled tolimit the instantaneous concentration of the first salt and the secondsalt in the aqueous mixture as desired. The first aqueous solutioncontaining the first salt and the second aqueous solution containing thesecond salt may also be dispersed directly into the third aqueoussolution at a flow rate selected to provide a desired instantaneousconcentration of the first salt and the second salt. The first aqueoussolution and the second aqueous solution may be dispersed directly intothe third aqueous solution using conventional means for dispersing onesolution into another solution at a controlled flow rate. For example,the first aqueous solution and the second aqueous solution may bedispersed into the third aqueous solution through separate nozzleslocated within the third aqueous solution, where the flow of the firstand second solutions through the nozzles is metered by separate flowmetering devices.

The particle size distribution of the catalyst material produced bymixing the first salt and the second salt in the aqueous mixture ispreferably controlled by the rate of addition of the first and secondaqueous solutions to the third aqueous solution, as described above, sothat the median and/or mean particle size of the particle sizedistribution falls within a range of from 50 nm to 5 μm. The particlesize distribution of the catalyst material may be controlled by the rateof addition of the first and second aqueous solutions to the thirdaqueous solution so that the median and/or mean particle size of theparticle size distribution of the catalyst material may range from atleast 50 nm up to 1 μm, or up to 750 μm, or up to 500 nm.

The surface area of the catalyst material particles produced by mixingthe first and second aqueous solutions in the third aqueous solution ispreferably controlled by the rate of addition of the first and secondaqueous solutions to the third aqueous solution, as described above, sothat the BET surface area of the catalyst material particles may rangefrom 50 m²/g to 500 m²/g. The surface area of the catalyst materialparticles may be controlled by the rate of addition of the first andsecond aqueous solutions to the third aqueous solution so that the BETsurface area of the catalyst material particles is from 100 m²/g to 350m²/g

The aqueous mixture containing the first salt and the second salt ismixed to facilitate interaction and reaction of the cationic componentof the first salt with the anionic component of the second salt to formthe catalyst material. The aqueous mixture may be mixed by anyconventional means for agitating an aqueous solution or an aqueousdispersion, for example by mechanical stirring.

During mixing of the aqueous mixture of the first and second salts, thetemperature of the aqueous mixture is maintained in the range of from15° C. to 150° C., or from 60° C. to 125° C., or from 65° C. to 100° C.When the cationic component of the second salt is ammonium, thetemperature should be maintained in a range from 65° C. to 150° C. toevolve ammonia as a gas from the second salt. The temperature of theaqueous mixture during mixing may be maintained at less than 100° C. sothat the mixing may be conducted without the application of positivepressure necessary to inhibit the water in the aqueous mixture frombecoming steam. If the second salt is a tetrathiostannate, thetemperature of the aqueous mixture may be maintained at 100° C. or lessto inhibit the degradation of the second salt into tin disulfides.

Maintaining the temperature of the aqueous mixture in a range of from50° C. to 150° C. may result in production of a catalyst material havinga relatively large surface area and a substantially reduced median ormean particle size relative to a catalyst material produced in the samemanner at a lower temperature. It is believed that maintaining thetemperature in the range of 50° C. to 150° C. drives the reaction of thecationic component of the first salt with the anionic component of thesecond salt, reducing the reaction time and limiting the time availablefor the resulting product to agglomerate prior to precipitation.Maintaining the temperature in a range of from 50° C. to 150° C. duringthe mixing of the first and second salts in the aqueous mixture mayresult in production of a catalyst material having a particle sizedistribution with a median or mean particle size of from 50 nm up to 5μm, or up to 1 μm, or up to 750 nm; and having a BET surface area offrom 50 m²/g up to 500 m²/g or from 100 m²/g to 350 m²/g.

The first and second salts in the aqueous mixture may be mixed under apressure of from 0.101 MPa to 10 MPa (1.01 bar to 100 bar). Preferably,the first and second salts in the aqueous mixture are mixed atatmospheric pressure, however, if the mixing is effected at atemperature greater than 100° C. the mixing may be conducted underpositive pressure to inhibit the formation of steam.

During mixing, the aqueous mixture of the first and second salts ismaintained under anaerobic conditions. Maintaining the aqueous mixtureunder anaerobic conditions during mixing inhibits the oxidation of thecatalyst material or the anionic component of the second salt so thatthe catalyst material produced by the process contains little, if anyoxygen other than oxygen present in the first and second salts. Theaqueous mixture of the first and second salts may be maintained underanaerobic conditions during mixing by conducting the mixing in anatmosphere containing little or no oxygen, preferably an inertatmosphere. The mixing of the first and second salts in the aqueousmixture may be conducted under nitrogen gas, argon gas, and/or steam tomaintain anaerobic conditions during the mixing. An inert gas,preferably nitrogen gas or steam, may be continuously injected into theaqueous mixture during mixing to maintain anaerobic conditions and tofacilitate mixing of the first and second salts in the aqueous mixtureand displacement of ammonia gas if the second salt contains an ammoniumcation.

The first and second salts may be mixed in the aqueous mixture at atemperature of from 15° C. to 150° C. under anaerobic conditions for aperiod of time sufficient to permit the formation of the preferredcatalyst material. The first and second salts may be mixed in theaqueous mixture for a period of at least 1 hour, or at least 2 hours, orat least 3 hours, or at least 4 hours, or from 1 hour to 10 hours, orfrom 2 hours to 9 hours, or from 3 hours to 8 hours, or from 4 hours to7 hours to form the catalyst material. The first and/or second salt(s)may be added to the aqueous mixture over a period of from 30 minutes to4 hours while mixing the aqueous mixture, and, after the entirety of thefirst and second salts have been mixed into the aqueous mixture, theaqueous mixture may be mixed for at least an additional 1 hour, or 2hours, or 3 hours or 4 hours, or 5 hours to form the catalyst material.

After completing mixing of the aqueous mixture of the first and secondsalts, a solid may be separated from the aqueous mixture to produce thepreferred catalyst material.

The solid may be separated from the aqueous mixture by any conventionalmeans for separating a solid phase material from a liquid phasematerial. For example, the solid may be separated by allowing the solidto settle from the resulting mixture, preferably for a period of from 1hour to 16 hours, and separating the solid from the mixture by vacuum orgravitational filtration or by centrifugation. To enhance recovery ofthe solid, water may be added to the aqueous mixture prior to allowingthe solid to settle. Water may be added to the aqueous mixture in avolume relative to the volume of the aqueous mixture of from 0.1:1 to0.75:1. Alternatively, but less preferably, the solid may be separatedfrom the mixture by centrifugation without first allowing the solid tosettle and/or without the addition of water. Alternatively, the aqueousmixture may be spray dried to separate the solid catalyst material fromthe aqueous mixture.

The preferred catalyst material may be washed subsequent to separationfrom the aqueous mixture, if desired. Substantial volumes of water maybe used to wash the separated catalyst material since the separatedcatalyst material is insoluble in water, and the yield of catalystmaterial will not be significantly affected by the wash.

Process for Cracking a Hydrocarbon-Containing Feedstock

In the process of the present invention, at least one catalyst asdescribed above, the hydrocarbon-containing feedstock, and hydrogen aremixed, preferably blended, at a temperature of from 375° C. to 500° C.and a total pressure of 6.9 MPa to 27.5 MPa. The hydrocarbon-containingfeedstock, the catalyst(s) and hydrogen may be mixed by contact witheach other in a mixing zone maintained at a temperature of from 375° C.to 500° C. and a total pressure of 6.9 MPa to 27.5 MPa, where thehydrocarbon-containing feedstock is continuously or intermittentlyprovided to the mixing zone at a rate of at least 350 kg/hr per m³ ofmixture volume in the mixing zone. A vapor that comprises hydrocarbonsthat are a gas at the temperature and pressure within the mixing zone isseparated from the mixing zone., Apart from the mixing zone, ahydrocarbon-containing product that comprises one or more hydrocarboncompounds that are liquid at STP may be condensed from the vaporseparated from the mixing zone. Alternatively, apart from the mixingzone the vapor separated from the mixing zone may be hydrotreated toreduce sulfur, nitrogen, and olefinic hydrocarbon content by contactingthe vapor with a commercially available hydrotreating catalyst andhydrogen at a temperature of from 260° C. to 425° C. and a totalpressure of from 3.4 MPa to 27.5 MPa.

In an embodiment of the process of the invention, as shown in FIG. 1,the mixing zone 1 may be in a reactor 3, where the conditions of thereactor 3 may be controlled to maintain the temperature and totalpressure in the mixing zone 1 at 375° C. to 500° C. and 6.9 MPa to 27.5MPa, respectively. The hydrocarbon-containing feedstock may be providedcontinuously or intermittently from a feed supply 2 to the mixing zone 1in the reactor 3 through feed inlet 5. The hydrocarbon-containingfeedstock may be preheated to a temperature of from 100° C. to 350° C.by a heating element 4, which may be a heat exchanger, prior to beingfed to the mixing zone 1.

The hydrocarbon-containing feedstock is provided to the mixing zone 1 ofthe reactor 3 at a rate of at least 350 kg/hr per m³ of the mixturevolume within mixing zone 1 of the reactor 3. The mixture volume isdefined herein as the combined volume of the catalyst, thehydrocarbon-depleted feed residuum (as defined herein), and thehydrocarbon-containing feedstock in the mixing zone 1, where thehydrocarbon-depleted feed residuum may contribute no volume to themixture volume (i.e. at the start of the process before ahydrocarbon-depleted feed residuum has been produced in the mixing zone1), and where the hydrocarbon-containing feedstock may contribute novolume to the mixture volume (i.e. after initiation of the processduring a period between intermittent addition of freshhydrocarbon-containing feedstock into the mixing zone 1). The mixturevolume within the mixing zone 1 may be affected by 1) the rate ofaddition of the hydrocarbon-containing feedstock into the mixing zone 1;2) the rate of removal of the vapor from the reactor 3; and, optionally,3) the rate at which a bleed stream of the hydrocarbon-depleted feedresiduum, catalyst, and hydrocarbon-containing feedstock is separatedfrom and recycled to the reactor 3, as described in further detailbelow. The hydrocarbon-containing feedstock may be provided to themixing zone 1 of the reactor 3 at a rate of at least 500, or at least600, or at least 700, or at least 800, or at least 900, or at least 1000kg/hr per m³ of the mixture volume within the mixing zone 1 up to 5000kg/hr per m³ of the mixture volume within the mixing zone 1.

Preferably, the mixture volume of the hydrocarbon-containing feedstock,the hydrocarbon-depleted feed residuum, and the catalyst is maintainedwithin the mixing zone within a selected range of the reactor volume byselecting 1) the rate at which the hydrocarbon-containing feedstock isprovided to the mixing zone 1; and/or 2) the rate at which a bleedstream is removed from and recycled to the mixing zone 1; and/or 3) thetemperature and pressure within the mixing zone 1 and the reactor 3 toprovide a selected rate of vapor removal from the mixing zone 1 and thereactor 3. The combined volume of the hydrocarbon-containing feedstockand the catalyst initially provided to the mixing zone 1 at the start ofthe process define an initial mixture volume, and the amount ofhydrocarbon-containing feedstock and the amount of the catalystinitially provided to the mixing zone 1 may be selected to provide aninitial mixture volume of from 5% to 97% of the reactor volume.,preferably from 30% to 75% of the reactor volume. The rate at which thehydrocarbon-containing feedstock is provided to the mixing zone 1 and/orthe rate at which a bleed stream is removed from and recycled to themixing zone 1 and/or the rate at which vapor is removed from the reactor3 and/or the temperature and total pressure within the mixing zone 1and/or the reactor 3 may be selected to maintain the mixture volume ofthe hydrocarbon-containing feedstock, the hydrocarbon-depleted feedresiduum, and the catalyst at a level of at least 10%, or at least 25%,or at least 40%, or at least 50%, or within 70%, or within 50%, or from10% to 1940%, or from 15% to 1000%, or from 20% to 500%, or from 25% to250%, or from 50% to 200% of the initial mixture volume during theprocess.

The hydrocarbon-containing feedstock may be provided to the mixing zone1 at such relatively high rates for reacting a feedstock containingrelatively large quantities of heavy, high molecular weight hydrocarbonsdue to the inhibition of coke formation in the process of the presentinvention. Conventional processes for cracking heavy hydrocarbonaceousfeedstocks are typically operated at rates on the order of 10 to 300kg/hr per m³ of reaction volume so that the conventional crackingprocess may be conducted either 1) at sufficiently low temperature toavoid excessive coke-make to maximize yield of desirable crackedhydrocarbons; or 2) at higher temperatures with significant quantitiesof coke production, where the high levels of solids produced impedesoperation of the process at a high rate.

Hydrogen is provided to the mixing zone 1 of the reactor 3 for mixing orblending with the hydrocarbon-containing feedstock and the catalyst.Hydrogen may be provided continuously or intermittently to the mixingzone 1 of the reactor 3 through hydrogen inlet line 7, or,alternatively, may be mixed together with the hydrocarbon-containingfeedstock, and optionally the catalyst, and provided to the mixing zone1 through the feed inlet 5. Hydrogen may be provided to the mixing zone1 of the reactor 3 at a rate sufficient to hydrogenate hydrocarbonscracked in the process. The hydrogen may be provided to the mixing zone1 in a ratio relative to the hydrocarbon-containing feedstock providedto the mixing zone 1 of from 1 Nm³/m³ to 16,100 Nm³/m³ (5.6 SCFB to90160 SCFB), or from 2 Nm³/m³ to 8000 Nm³/m³ (11.2 SCFB to 44800 SCFB),or from 3 Nm³/m³ to 4000 Nm³/m³ (16.8 SCFB to 22400 SCFB), or from 5Nm³/m³ to 320 Nm³/m³ (28 SCFB to 1792 SCFB). The hydrogen partialpressure in the mixing zone 1 may be maintained in a pressure range offrom 2.1 MPa to 27.5 MPa, or from 5 MPa to 20 MPa, or from 10 MPa to 15MPa.

The catalyst may be located in the mixing zone 1 in the reactor 3 or maybe provided to the mixing zone 1 in the reactor 3 during the process ofthe present invention. Catalysts that may be utilized in the process areas described above, and exclude catalysts exhibiting significant acidityincluding catalysts having an acidity as measured by ammoniachemisorption of more than 200 mmol ammonia per gram of catalyst. Thecatalyst may be located in the mixing zone 1 in a catalyst bed.Preferably, however, the catalyst is provided to the mixing zone 1during the process, or, if located in the mixing zone initially, may beblended with the hydrocarbon-containing feed and hydrogen, and is notpresent in a catalyst bed. The catalyst may be provided to the mixingzone 1 together with the hydrocarbon-containing feedstock through feedinlet 5, where the catalyst may be dispersed in thehydrocarbon-containing feedstock prior to feeding the mixture to themixing zone 1 through the feed inlet 5. Alternatively, the catalyst maybe provided to the mixing zone 1 through a catalyst inlet 9, where thecatalyst may be mixed with sufficient hydrocarbon-containing feedstockor another fluid, for example a hydrocarbon-containing fluid, to enablethe catalyst to be delivered to the mixing zone 1 through the catalystinlet 9.

The catalyst is provided to be mixed with the hydrocarbon-containingfeedstock and the hydrogen in the mixing zone 1 in a sufficient amountto catalytically crack the hydrocarbon-containing feedstock and/or tocatalyze hydrogenation of the cracked hydrocarbons in the mixing zone.An initial charge of the catalyst may be provided for mixing with aninitial charge of hydrocarbon-containing feedstock in an amount of from20 g to 125 g of catalyst per kg of initial hydrocarbon-containingfeedstock. Over the course of the process, the catalyst may be providedfor mixing with the hydrocarbon-containing feedstock and hydrogen in anamount of from 0.125 g to 5 g of catalyst per kg ofhydrocarbon-containing feedstock. Alternatively, the catalyst may beprovided for mixing with the hydrocarbon-containing feedstock andhydrogen over the course of the process in an amount of from 0.125 g to50 g of catalyst per kg of hydrocarbons in the hydrocarbon-containingfeedstock having a boiling point of at least 538° C. at a pressure of0.101 MPa.

The catalyst, the hydrocarbon-containing feedstock, and the hydrogen maybe mixed by being blended into an intimate admixture in the mixing zone1. The catalyst, hydrocarbon-containing feedstock and the hydrogen maybe blended in the mixing zone 1, for example, by stirring a mixture ofthe components, for example by a mechanical stirring device located inthe mixing zone 1. The catalyst, hydrocarbon-containing feedstock, andhydrogen may also be mixed in the mixing zone 1 by blending thecomponents prior to providing the components to the mixing zone 1 andinjecting the blended components into the mixing zone 1 through one ormore nozzles which may act as the feed inlet 5. The catalyst,hydrocarbon-containing feedstock, and hydrogen may also be blended inthe mixing zone 1 by blending the hydrocarbon-containing feedstock andcatalyst and injecting the mixture into the mixing zone 1 through one ormore feed inlet nozzles positioned with respect to the hydrogen inletline 7 such that the mixture is blended with hydrogen entering themixing zone 1 through the hydrogen inlet line 7. Baffles may be includedin the reactor 3 in the mixing zone 1 to facilitate blending thehydrocarbon-containing feedstock, catalyst, and hydrogen. Lesspreferably, the catalyst is present in the mixing zone 1 in a catalystbed, and the hydrocarbon-containing feedstock, hydrogen, and catalystare mixed by bringing the hydrocarbon-containing feedstock and hydrogensimultaneously into contact with the catalyst in the catalyst bed.

The temperature and pressure conditions in the mixing zone 1 aremaintained so that heavy hydrocarbons in the hydrocarbon-containingfeedstock may be cracked. The temperature in the mixing zone 1 ismaintained from 375° C. to 500° C. Preferably, the mixing zone 1 ismaintained at a temperature of from 425° C. to 500° C., or from 430° C.to 500° C., or from 440° C. to 500° C., or from 450° C. to 500° C. In anembodiment of the process of the present invention, the temperaturewithin the mixing zone is selected and controlled to be at least 430°C., or at least 450° C. Higher temperatures may be preferred in theprocess of the present invention since 1) the rate of conversion of thehydrocarbon-containing feedstock to a hydrocarbon-containing productsignificantly increases with temperature; and 2) the present processinhibits or prevents the formation of coke, even at temperatures of 430°C. or greater, or 450° C. or greater, which typically occurs rapidly inconventional cracking processes at temperatures of 430° C. or greater,or 450° C. or greater.

Mixing the hydrocarbon-containing feedstock, the catalyst(s), andhydrogen in the mixing zone 1 at a temperature of from 375° C. to 500°C. and a total pressure of from 6.9 MPa to 27.5 MPa produces a vaporcomprised of hydrocarbons that are vaporizable at the temperature andpressure within the mixing zone 1. The vapor may be comprised ofhydrocarbons present initially in the hydrocarbon-containing feedstockthat vaporize at the temperature and pressure within the mixing zone 1and hydrocarbons that are not present initially in thehydrocarbon-containing feedstock but are produced by cracking andhydrogenating hydrocarbons initially in the hydrocarbon-containingfeedstock that were not vaporizable at the temperature and pressurewithin the mixing zone 1 prior to cracking.

At least a portion of the vapor comprised of hydrocarbons that arevaporizable at the temperature and pressure within the mixing zone 1 maybe continuously or intermittently separated from the mixing zone 1containing the mixture of hydrocarbon-containing feedstock, hydrogen,and catalyst since the more volatile vapor physically separates from thehydrocarbon-containing feedstock, catalyst, and hydrogen mixture. Thevapor may also contain hydrogen gas and hydrogen sulfide gas, which alsoseparate from the mixture in the mixing zone 1.

Separation of the vapor from the mixture in the mixing zone 1 leaves ahydrocarbon-depleted feed residuum from which the hydrocarbons presentin the vapor have been removed. The hydrocarbon-depleted feed residuumis comprised of hydrocarbons that are liquid at the temperature andpressure within the mixing zone 1. The hydrocarbon-depleted feedresiduum may also be comprised of solids such as metals freed fromcracked hydrocarbons and minor amounts of coke. The hydrocarbon-depletedfeed residuum may contain little coke or proto-coke since the process ofthe present invention inhibits the generation of coke. Thehydrocarbon-depleted feed residuum may contain, per metric ton ofhydrocarbon feedstock provided to the mixing zone 1, less than 30 kg, orat most 20 kg, or at most 10 kg, or at most 5 kg of hydrocarbonsinsoluble in toluene as measured by ASTM Method D4072.

At least a portion of the hydrocarbon-depleted feed residuum is retainedin the mixing zone 1 while the vapor is separated from the mixing zone1. The portion of the hydrocarbon-depleted feed residuum retained in themixing zone 1 may be subject to further cracking to produce more vaporthat may be separated from the mixing zone 1 and then from the reactor 3from which the liquid hydrocarbon-containing product may be produced bycooling. Hydrocarbon-containing feedstock and hydrogen may becontinuously or intermittently provided to the mixing zone 1 at therates described above and mixed with the catalyst and thehydrocarbon-depleted feed residuum retained in the mixing zone 1 toproduce further vapor comprised of hydrocarbons that are vaporizable atthe temperature and pressure within the mixing zone 1 for separationfrom the mixing zone 1 and the reactor 3.

At least a portion of the vapor separated from the mixture of thehydrocarbon-containing feedstock, hydrogen, and catalyst may becontinuously or intermittently separated from the mixing zone 1 whileretaining the hydrocarbon-depleted feed residuum, catalyst, and anyfresh hydrocarbon-containing feedstock in the mixing zone 1. At least aportion of the vapor separated from the mixing zone 1 may becontinuously or intermittently separated from the reactor 3 through areactor product outlet 11. The reactor 3 is preferably configured andoperated so that substantially only vapors and gases may exit thereactor product outlet 11, where the vapor product exiting the reactor 3comprises at most 5 wt. %, or at most 3 wt. %, or at most 1 wt. %, or atmost 0.5 wt. %, or at most 0.1 wt. %, or at most 0.01 wt. %, or at most0.001 wt. % solids and liquids at the temperature and pressure at whichthe vapor product exits the reactor 3.

A stripping gas may be injected into the reactor 3 over the mixing zone1 to facilitate separation of the vapor from the mixing zone 1. Thestripping gas may be heated to a temperature at or above the temperaturewithin the mixing zone 1 to assist in separating the vapor from themixing zone 1. In an embodiment of the process, the stripping gas may behydrogen gas and/or hydrogen sulfide gas.

As shown in FIG. 2, the reactor 3 may be comprised of a mixing zone 1, adisengagement zone 21, and a vapor/gas zone 23. The vapor comprised ofhydrocarbons that are vaporizable at the temperature and pressure withinthe mixing zone 1 may separate from the mixture of hydrocarbon-depletedresiduum, catalyst, hydrogen, and fresh hydrocarbon-containing feed, ifany, in mixing zone 1 into the disengagement zone 21. A stripping gassuch as hydrogen may be injected into the disengagement zone 21 tofacilitate separation of the vapor from the mixing zone 1. Some liquidsand solids may be entrained by the vapor as it is separated from themixing zone 1 into the disengagement zone 21, so that the disengagementzone 21 contains a mixture of vapor and liquids, and potentially solids.At least a portion of the vapor separates from the disengagement zone 21into the vapor/gas zone 23, where the vapor separating from thedisengagement zone 21 into the vapor/gas zone 23 contains little or noliquids or solids at the temperature and pressure within the vapor/gaszone. At least a portion of the vapor in the vapor/gas zone 23 exits thereactor 3 through the reactor product outlet 11.

Referring now to FIGS. 1 and 2, in the process of the present invention,the hydrocarbons in the hydrocarbon-containing feed andhydrocarbon-containing feed residuum are contacted and mixed with thecatalyst and hydrogen in the mixing zone 1 of the reactor 3 only as longas necessary to be vaporized and separated from the mixture, and areretained in the reactor 3 only as long as necessary to be vaporized andexit the reactor product outlet 11. Low molecular weight hydrocarbonshaving a low boiling point may be vaporized almost immediately uponbeing introduced into the mixing zone 1 when the mixing zone 1 ismaintained at a temperature of 375° C. to 500° C. and a total pressureof from 6.9 MPa to 27.5 MPa. These hydrocarbons may be separated rapidlyfrom the reactor 3. High molecular weight hydrocarbons having a highboiling point, for example hydrocarbons having a boiling point greaterthan 538° C. at 0.101 MPa, may remain in the mixing zone 1 until theyare cracked and hydrogenated into hydrocarbons having a boiling pointlow enough to be vaporized at the temperature and pressure in the mixingzone 1 and to exit the reactor 3. The hydrocarbons of thehydrocarbon-containing feed, therefore, are contacted and mixed with thecatalyst and hydrogen in the mixing zone 1 of the reactor 3 for avariable time period, depending on the boiling point of the hydrocarbonsunder the conditions in the mixing zone 1 and the reactor 3.

The rate of the process of producing the vapor product from thehydrocarbon-containing feedstock may be adjusted by selection of thetemperature and/or total pressure in the reactor 3, and particularly inthe mixing zone 1, within the temperature range of 375° C.-500° C. andwithin the pressure range of 6.9 MPa-27.5 MPa. Increasing thetemperature and/or decreasing the pressure in the mixing zone 1 permitsthe hydrocarbon-containing feedstock to provided to the reactor 3 at anincreased rate and the vapor product to be removed from the reactor 3 atan increased rate since the hydrocarbons in the hydrocarbon-containingfeedstock may experience a decreased residence time in the reactor 3 dueto higher cracking activity and/or faster vapor removal. Conversely,decreasing the temperature and/or increasing the pressure in the mixingzone 1 may reduce the rate at which the hydrocarbon-containing feedstockmay be provided to the reactor 3 and the vapor product may be removedfrom the reactor 3 since the hydrocarbons in the hydrocarbon-containingfeedstock may experience an increased residence time in the reactor 3due to lower cracking activity and/or slower vapor removal.

As a result of the inhibition and/or prevention of the formation of cokein the process, the hydrocarbons in the hydrocarbon-containing feed maybe contacted and mixed with the catalyst and hydrogen in the mixing zone1 at a temperature of 375° C. to 500° C. and a total pressure of 6.9 MPato 27.5 MPa for as long as necessary to be vaporized, or to be cracked,hydrogenated, and vaporized. It is believed that high boiling, highmolecular weight hydrocarbons may remain in the mixing zone 1 in thepresence of cracked hydrocarbons since the catalyst promotes theformation of hydrocarbon radical anions upon cracking that react withhydrogen to form stable hydrocarbon products rather than hydrocarbonradical cations that react with other hydrocarbons to form coke. Cokeformation is also avoided because the cracked hydrogenated hydrocarbonspreferentially exit the mixing zone 1 as a vapor rather remaining in themixing zone 1 to combine with hydrocarbon radicals in the mixing zone 1to form coke or proto-coke.

At least a portion of the vapor separated from the mixing zone 1 andseparated from the reactor 3 may be condensed apart from the mixing zone1 to produce a liquid hydrocarbon-containing product, Referring now toFIG. 1, the portion of the vapor separated from the reactor 3 may beprovided to a condenser 13 wherein at least a portion of the vaporseparated from the reactor 3 may be condensed to produce thehydrocarbon-containing product that is comprised of hydrocarbons thatare liquid at STP. A portion of the vapor separated from the reactor 3may be passed through a heat exchanger 15 to cool the vapor prior toproviding the vapor to the condenser 13.

Condensation of the liquid hydrocarbon-containing product from the vaporseparated from the reactor 3 may also produce a non-condensable gas thatmay be comprised of hydrocarbons having a carbon number from 1 to 6,hydrogen, and hydrogen sulfide. The condensed hydrocarbon-containingliquid product may be separated from the non-condensable gas through acondenser liquid product outlet 17 and stored in a product receiver 18,and the non-condensable gas may be separated from the condenser 13through a non-condensable gas outlet 19 and passed through an amine orcaustic scrubber 20 and recovered through a gas product outlet 22.

Alternatively, referring now to FIG. 2, the portion of the vaporseparated from the reactor 3 may be provided to a high pressureseparator 12 to separate a liquid hydrocarbon-containing product fromgases not condensable at the temperature and pressure within the highpressure separator 12, and the liquid hydrocarbon-containing productcollected from the high pressure separator may be provided through line16 to a low pressure separator 14 operated at a pressure less than thehigh pressure separator 12 to separate the liquid hydrocarbon-containingproduct from gases that are not condensable at the temperature andpressure at which the low pressure separator 14 is operated. Thevapor/gas exiting the reactor 3 from the reactor product outlet 11 maybe cooled prior to being provided to the high pressure separator 12 bypassing the vapor/gas through heat exchanger 15. The condensedhydrocarbon-containing liquid product may be separated from thenon-condensable gas in the low pressure separator through a low pressureseparator liquid product outlet 10 and stored in a product receiver 18.The non-condensable gas may be separated from the high pressureseparator 12 through a high pressure non-condensable gas outlet 24 andfrom the low pressure separator 14 through a low pressurenon-condensable gas outlet 26. The non-condensable gas streams may becombined in line 28 and passed through an amine or caustic scrubber 20and recovered through a gas product outlet 22.

Alternatively, the vapor separated from the mixing zone 1 and from thereactor 3 may be further hydroprocessed without condensing thehydrocarbon-containing product from the vapor. For example, the vaporseparated from the reactor may be hydrotreated to reduce sulfur,nitrogen, and olefins in the hydrocarbon-containing product by passingthe vapor from the reactor 3 to a hydroprocessing reactor, where thevapor may be contacted with a conventional hydroprocessing catalyst andhydrogen at a temperature of from 260° C. to 425° C. and a totalpressure of from 3.4 MPa to 27.5 MPa.

A portion of the hydrocarbon-depleted feed residuum and catalyst may beseparated from the mixing zone to remove solids including metals andhydrocarbonaceous solids including coke from the hydrocarbon-depletedfeed residuum Referring now to FIGS. 1 and 2, the reactor 3 may includea bleed stream outlet 25 for removal of a stream of hydrocarbon-depletedfeed residuum and catalyst from the mixing zone 1 and the reactor 3. Thebleed stream outlet 25 may be operatively connected to the mixing zone 1of the reactor 3.

A portion of the hydrocarbon-depleted feed residuum and the catalyst maybe removed together from the mixing zone 1 and the reactor 3 through thebleed stream outlet 25 while the process is proceeding. Solids and thecatalyst may be separated from a liquid portion of thehydrocarbon-depleted feed residuum in a solid-liquid separator 30. Thesolid-liquid separator 30 may be a filter or a centrifuge. The liquidportion of the hydrocarbon-depleted feed residuum may be recycled backinto the mixing zone 1 via a recycle inlet 32 for further processing ormay be combined with the hydrocarbon-containing feed and recycled intothe mixing zone 1 through the feed inlet 5.

In a preferred embodiment, hydrogen sulfide is mixed, and preferablyblended, with the hydrocarbon-containing feedstock, hydrogen, anyhydrocarbon-depleted feed residuum, and the catalyst in the mixing zone1 of the reactor 3. The hydrogen sulfide may be provided continuously orintermittently to the mixing zone 1 of the reactor 3 as a liquid or agas. The hydrogen sulfide may be mixed with the hydrocarbon-containingfeedstock and provided to the mixing zone 1 with thehydrocarbon-containing feedstock through the feed inlet 5.Alternatively, the hydrogen sulfide may be mixed with hydrogen andprovided to the mixing zone 1 through the hydrogen inlet line 7.Alternatively, the hydrogen sulfide may be provided to the mixing zone 1through a hydrogen sulfide inlet line 27.

It is believed that hydrogen sulfide acts as a further catalyst incracking hydrocarbons in the hydrocarbon-containing feedstock in thepresence of hydrogen and the catalyst and lowers the activation energyto crack hydrocarbons in the hydrocarbon-containing feed stock, therebyincreasing the rate of the reaction. The rate of the process, inparticular the rate that the hydrocarbon-containing feedstock may beprovided to the mixing zone 1 for cracking and cracked product may beremoved from the reactor 3, therefore, may be greatly increased with theuse of significant quantities of hydrogen sulfide in the process. Forexample, the rate of the process may be increased by at least 1.5 times,or by at least 2 times, the rate of the process in the absence ofsignificant quantities of hydrogen sulfide.

As discussed above, it is also believed that the hydrogen sulfide actingas a further catalyst inhibits coke formation under cracking conditions.Use of sufficient hydrogen sulfide in the process permits the process tobe effected at a mixing zone temperature of at least at least 430° C. orat least 450° C. with little or no increase in coke formation relativeto cracking conducted at lower temperatures since hydrogen sulfideinhibits coke formation. The rate of the process, in particular the ratethat the hydrocarbon-containing feedstock may be provided to the mixingzone 1 for cracking and cracked product may be removed from the reactor3, therefore, may be greatly increased with the use of significantquantities of hydrogen sulfide in the process since the rate of reactionin the process increases significantly relative to temperature, and thereaction may be conducted at higher temperatures in the presence ofhydrogen sulfide without significant coke production.

The hydrogen sulfide provided to be mixed with thehydrocarbon-containing feedstock, hydrogen, and the catalyst may beprovided in an amount effective to increase the rate of the crackingreaction. In order to increase the rate of the cracking reaction,hydrogen sulfide may be provided in an amount on a mole ratio basisrelative to hydrogen provided to be mixed with thehydrocarbon-containing feedstock and catalyst, of at least 0.5 mole ofhydrogen sulfide per 9.5 moles hydrogen, where the combined hydrogensulfide and hydrogen partial pressures are maintained to provide atleast 60%, or at least 70%, or at least 80%, or at least 90%, or atleast 95% of the total pressure in the reactor. The hydrogen sulfide maybe provided in an amount on a mole ratio basis relative to the hydrogenprovided of at least 1:9, or at least 1.5:8.5, or at least 2.5:7.5, orat least 3:7 or at least 3.5:6.5, or at least 4:6, up to 1:1, where thecombined hydrogen sulfide and hydrogen partial pressures are maintainedto provide at least 60%, or at least 70%, or at least 80%, or at least90%, or at least 95% of the total pressure in the reactor. The hydrogensulfide partial pressure in the reactor may be maintained in a pressurerange of from 0.4 MPa to 13.8 MPa, or from 2 MPa to 10 MPa, or from 3MPa to 7 MPa.

The combined partial pressure of the hydrogen sulfide and hydrogen inthe reactor may be maintained to provide at least 60% of the totalpressure in the reactor, where the hydrogen sulfide partial pressure ismaintained at a level of at least 5% of the hydrogen partial pressure.Preferably, the combined partial pressure of the hydrogen sulfide andhydrogen in the reactor is maintained to provide at least 70%, or atleast 75%, or at least 80%, or at least 90%, or at least 95% of thetotal pressure in the reactor, where the hydrogen sulfide partialpressure is maintained at a level of at least 5% of the hydrogen partialpressure. Other gases may be present in the reactor in minor amountsthat provide a pressure contributing to the total pressure in thereactor. For example, a non-condensable gas produced in the vapor alongwith the hydrocarbon-containing product may be separated from thehydrocarbon-containing product and recycled back into the mixing zone,where the non-condensable gas may comprise hydrocarbon gases such asmethane, ethane, and propane as well as hydrogen sulfide and hydrogen.

The vapor separated from the mixing zone 1 and from the reactor 3through the reactor product outlet 11 may contain hydrogen sulfide. Thehydrogen sulfide in the vapor product may be separated from thehydrocarbon-containing liquid product in the condenser 13 (FIG. 1) or inthe high and low pressure separators 12 and 14 (FIG. 2), where thehydrogen sulfide may form a portion of the non-condensable gas. Whenhydrogen sulfide is provided to the mixing zone 1 in the process, it ispreferable to condense the hydrocarbon-containing liquid product at atemperature of from 60° C. to 93° C. (140° F.-200° F.) so that hydrogensulfide is separated from the hydrocarbon-containing liquid product withthe non-condensable gas rather than condensing with the liquidhydrocarbon-containing product. The non-condensable gas including thehydrogen sulfide may be recovered from the condenser 13 through the gasproduct outlet 19 (FIG. 1) or from the high pressure separator 12through high pressure separator gas outlet 24 and the low pressureseparator gas outlet 26 (FIG. 2). The hydrogen sulfide may be separatedfrom the other components of the non-condensable gas by treatment of thenon-condensable gas to recover the hydrogen sulfide. For example, thenon-condensable gas may be scrubbed with an amine solution in thescrubber 20 to separate the hydrogen sulfide from the other componentsof the non-condensable gas. The hydrogen sulfide may then be recoveredand recycled back into the mixing zone 1.

Alternatively, the vapor containing hydrogen sulfide may behydroprocessed as described above by contacting the vapor with ahydroprocessing catalyst and hydrogen at a temperature of from 260° C.to 425° C. and a total pressure of from 3.4 MPa to 27.5 MPa withoutfirst condensing a liquid hydrocarbon-containing product. Thehydrotreated vapor may contain hydrocarbons that are liquid at STP thatmay be condensed and separated from non-condensable hydrocarbons,hydrogen, and hydrogen sulfide. The non-condensable hydrocarbons,hydrogen, and hydrogen sulfide may be recycled into the mixing zone, orthe hydrogen sulfide may be separated from the non-condensablehydrocarbons and hydrogen by scrubbing with an amine solution, where theseparated hydrogen sulfide may be regenerated from the amine solutionand recycled to the mixing zone.

The process of the present invention may be effected for a substantialperiod of time on a continuous or semi-continuous basis, in part becausethe process generates little or no coke. The hydrocarbon-containingfeedstock, hydrogen, catalyst, and hydrogen sulfide (if used in theprocess) may be continuously or intermittently provided to the mixingzone 1 in the reactor 3, where the hydrocarbon-containing feedstock isprovided at a rate of at least 350 kg/hr per m³ of the mixture volume asdefined above, and mixed in the mixing zone 1 at a temperature of from375° C.-500° C. and a total pressure of from 6.9 MPa-27.5 MPa for aperiod of at least 40 hours, or at least 100 hours, or at least 250hours, or at least 500 hours, or at least 750 hours to generate thevapor comprised of hydrocarbons that are vaporizable at the temperatureand pressure in the mixing zone 1 and the hydrocarbon-depleted feedresiduum, as described above. The vapor may be continuously orintermittently separated from the mixing zone 1 and the reactor 3 oversubstantially all of the time period that the hydrocarbon-containingfeedstock, catalyst, hydrogen, and hydrogen sulfide, if any, are mixedin the mixing zone 1. Fresh hydrocarbon-containing feedstock, hydrogen,and hydrogen sulfide, if used in the process, may be blended with thehydrocarbon-depleted feed residuum and catalyst in the mixing zone 1over the course of the time period of the reaction as needed. In apreferred embodiment, fresh hydrocarbon-containing feedstock, hydrogen,and hydrogen sulfide, if any, are provided continuously to the mixingzone 1 over substantially all of the time period the reaction iseffected. Solids may be removed from the mixing zone 1 continuously orintermittently over the time period the process is run by separating ableed stream of the hydrocarbon-containing feed residuum from the mixingzone 1 and the reactor 3, removing the solids from the bleed stream, andrecycling the bleed stream from which the solids have been removed backinto the mixing zone 1 as described above.

The process of the present invention may produce, in part, ahydrocarbon-containing product that is a liquid at STP. Thehydrocarbon-containing product may contain less than 4 wt. %, or lessthan 3 wt. %, or at most 2 wt. %, or at most 1 wt. %, or at most 0.5 wt.% of hydrocarbons having a boiling point of greater than 538° C. asdetermined in accordance with ASTM Method D5307 and may contain at most0.5 wt. % or at most 0.25 wt. %, or at most 0.1 wt. % coke as determinedin accordance with ASTM Method D4072. Furthermore, thehydrocarbon-containing product may contain at least 80%, or at least85%, or at least 90%, or at least 95%, or at least 97% of the atomiccarbon present in the hydrocarbon-containing feedstock. Therefore, whenthe process of the present invention is utilized, most of thehydrocarbons in the hydrocarbon-containing feedstock may be recovered inthe hydrocarbon-containing product that is liquid at STP, and little ofthe hydrocarbons in the hydrocarbon-containing feedstock are convertedto coke or gas.

The hydrocarbon-containing product may contain VGO hydrocarbons,distillate hydrocarbons, and naphtha hydrocarbons. Thehydrocarbon-containing product may contain, per gram, at least 0.05grams, or at least 0.1 grams of hydrocarbons having a boiling point fromthe initial boiling point of the hydrocarbon-containing product up to204° C. (400° F.). The hydrocarbon-containing product may also contain,per gram, at least 0.1 grams, or at least 0.15 grams of hydrocarbonshaving a boiling point of from 204° C. (400° F.) up to 260° C. (500°F.). The hydrocarbon-containing product may also contain, per gram, atleast 0.25 grams, or at least 0.3 grams, or at least 0.35 grams ofhydrocarbons having a boiling point of from 260° C. (500° F.) up to 343°C. (650° F.). The hydrocarbon-containing product may also contain, pergram, at least 0.3 grams, or at least 0.35 grams, or at least 0.4, or atleast 0.45 grams of hydrocarbons having a boiling point of from 343° C.(500° F.) up to 538° C. (1000° F.). The relative amounts of hydrocarbonswithin each boiling range and the boiling range distribution of thehydrocarbons may be determined in accordance with ASTM Method D5307.

The hydrocarbon-containing product produced by the process of thepresent invention may contain significant amounts of sulfur, providedthe hydrocarbon-containing product is condensed from the vapor separatedfrom the mixing zone without first hydroprocessing the vapor. Thehydrocarbon-containing product may contain, per gram, at least 0.0005gram of sulfur or at least 0.001 gram of sulfur. The sulfur content ofthe hydrocarbon-containing product may be determined in accordance withASTM Method D4294. At least 40 wt. % of the sulfur may be contained inhydrocarbon compounds having a carbon number of 17 or less as determinedby two-dimensional GC-GC sulfur chemiluminscence, where at least 60 wt.% of the sulfur in the sulfur-containing hydrocarbon compounds having acarbon number of 17 or less may be contained in benzothiopheniccompounds as determined by GC-GC sulfur chemiluminscence.

The hydrocarbon-containing product produced by the process of thepresent invention may contain significant amounts of nitrogen, providedthe hydrocarbon-containing product is condensed from the vapor separatedfrom the mixing zone without first hydroprocessing the vapor. Thehydrocarbon-containing product produced by the process of the presentinvention may contain, per gram, at least 0.0005 gram or at least 0.001gram of nitrogen as determined in accordance with ASTM Method D5762. Thehydrocarbon-containing product may have a relatively low ratio of basicnitrogen compounds to other nitrogen containing compounds therein. Thenitrogen may be contained in hydrocarbon compounds, where at least 30wt. % of the nitrogen in the hydrocarbon composition is contained innitrogen-containing hydrocarbon compounds having a carbon number of 17or less and where at least 50 wt. % of the nitrogen-containinghydrocarbon compounds having a carbon number of 17 or less are acridinicand carbazolic compounds. The amount of nitrogen-containing hydrocarboncompounds having a carbon number of 17 or less relative to the amount ofnitrogen in all nitrogen-containing hydrocarbon compounds in thehydrocarbon-containing product and the relative amount of acridinic andcarbazolic compounds may be determined by nitrogen chemiluminscence twodimensional gas chromatography (GC×GC—NCD).

The hydrocarbon-containing product produced by the process of thepresent invention may contain significant quantities of aromatichydrocarbon compounds. The hydrocarbon-containing product may contain,per gram, at least 0.3 gram, or at least 0.35 gram, or at least 0.4gram, or at least 0.45 gram, or at least 0.5 gram of aromatichydrocarbon compounds.

The hydrocarbon-containing product of the process of the presentinvention may contain relatively few polyaromatic hydrocarbon compoundscontaining three or more aromatic ring structures (e.g. anthracene,phenanthrene) relative to mono-aromatic and di-aromatic hydrocarboncompounds. The combined mono-aromatic and di-aromatic hydrocarboncompounds in the hydrocarbon-containing product may be present in thehydrocarbon-containing product in a weight ratio relative to thepolyaromatic hydrocarbon compounds (containing three or more aromaticring structures) of at least 1.5:1.0, or at least 2.0:1.0, or at least2.5:1.0. The relative amounts of mono-aromatic, di-aromatic, andpolyaromatic compounds in the hydrocarbon-containing product may bedetermined by flame ionization detection-two dimensional gaschromatography (GC×GC—FID).

To facilitate a better understanding of the present invention, thefollowing examples of certain aspects of some embodiments are given. Inno way should the following examples be read to limit, or define, thescope of the invention.

Example 1

A catalyst for use in a process of the present invention containingcopper, molybdenum, and sulfur was produced, where at least a portion ofthe catalyst had a structure according to Formula (XVII):

781 grams of ammonium tetrathiomolybdate was mixed with 636 grams ofNa₂CO₃ in 6 liters of water while stirring. The resulting solution washeated to 70° C. and then stirred for three hours to produce a solutionof Na₂MoS₄. The Na₂MoS₄ solution was then permitted to cool overnight. Asecond solution was prepared by mixing 1498 grams of CuSO₄5H₂O in 6liters of water. The CuSO₄ solution was then added to the Na₂MoS₄solution via pneumatic pump through a 0.02″×0.5″ nozzle while stirringthe mixture at ambient temperature. The mixture was stirred for twohours, and then the resulting solids were separated by centrifuge. 880grams of solid particulate catalyst was recovered. The solids were thenwashed with water until the effluent from the wash had a conductivity of488 μS at 33° C. The catalyst solids were particulate and had a particlesize distribution with a mean particle size of 8.5 μm as determined bylaser diffractometry using a Mastersizer S (Malvern Instruments). TheBET surface area of the catalyst solids was measured to be 29.3 m²/g.Semi-quantitative XRF of the catalyst solids indicated that the catalystsolids contained, by mass, 45.867% Cu, 18.587% Mo, and 27.527% S. X-raydiffraction and Raman IR spectroscopy confirmed that at least a portionof the catalyst had a structure in which copper, molybdenum, and sulfurwere arranged as shown in formula (XVII) above.

A solid material prepared in a similar manner was determined to have anacidity of 70 μmol of ammonia uptake per gram of solid material.

Example 2

Bitumen from Peace River, Canada was selected as ahydrocarbon-containing feedstock for cracking. The Peace River bitumenwas analyzed to determine its composition. The properties of the PeaceRiver bitumen are set forth in Table 1:

TABLE 1 Property Value Hydrogen (wt. %) 10.1 Carbon (wt. %) 82 Oxygen(wt. %) 0.62 Nitrogen (wt. %) 0.37 Sulfur (wt. %) 6.69 Nickel (wppm) 70Vanadium (wppm) 205 Microcarbon residue (wt. %) 12.5 C5 asphaltenes (wt.%) 10.9 Density (g/ml) 1.01 Viscosity at 38° C. (cSt) 8357 TAN-E (ASTMD664) (mg KOH/g) 3.91 Boiling Range Distribution Initial BoilingPoint-204° C. (400° F.)(wt. %) [Naphtha] 0 204° C. (400° F.)-260° C.(500° F.) (wt. %) [Kerosene] 1 260° C. (500° F.)-343° C. (650° F.) (wt.%) [Diesel] 14 343° C. (650° F.)-538° C. (1000° F.) (wt. %) [VGO]37.5 >538° C. (1000° F.) (wt. %) [Residue] 47.5

Peace River bitumen having the composition shown in Table 1 above washydrocracked in a process in accordance with the present invention usingdifferent hydrogen sulfide and hydrogen levels. Hydrogen sulfide wasprovided at 0 mol %, 5 mol %, 11.4 mol %, and 20.1 mol % of the gas fedto the reactor. Hydrogen was provided at: 70.2 mol % of the gas fed tothe reactor when 0 mol % hydrogen sulfide was fed to the reactor; 70 mol% of the gas fed to the reactor when hydrogen sulfide was provided at 5mol % (mole ratio of 1:14, hydrogen sulfide:hydrogen); 68.6 mol % of thegas fed to the reactor when hydrogen sulfide was provided at 11.4 mol %(mole ratio of 1:6, hydrogen sulfide:hydrogen); and 69.9 mol % of thegas fed to the reactor when hydrogen sulfide was provided at 20.1 mol %(mole ratio of 1:3.5, hydrogen sulfide:hydrogen). Nitrogen was providedas an inert gas in the gas fed to the reactor to maintain the totalpressure of the reaction at 8.3 MPa, where nitrogen was provided as 25mol % of the gas fed to the reactor when hydrogen sulfide was providedat 5 mol %; as 20 mol % of the gas fed to the reactor when hydrogensulfide was provided at 11.4 mol %; as 10 mol % of the gas fed to thereactor when hydrogen sulfide was provided at 20.1 mol %; and as 29.8mol % of the gas fed to the reactor when the gas fed to the reactorcontained no hydrogen sulfide and 70.2 mol % hydrogen. Hydrogen andhydrogen sulfide provided 75% of the total pressure in the reaction whenhydrogen sulfide was provided at 5 mol % of the gas fed to the reactor,and provided 80% of the total pressure when hydrogen sulfide wasprovided at 11.4 mol % and 20.1 mol % of the gas fed to the reactor.Hydrogen provided 70.2% of the total pressure when only hydrogen andnitrogen were provided to the reactor.

Four samples of the Peace River bitumen were hydrocracked, one each atthe above specified hydrogen sulfide: hydrogen: nitrogen levels. Thetotal pressure of each hydrocracking reaction was maintained at 8.3 MPaand the temperature of each hydrocracking reaction was maintained at430° C. The hydrogen, hydrogen sulfide and nitrogen gases were providedtogether to each hydrocracking reaction at a gas flow rate of 900standard liters per hour. In the hydrocracking treatment of each sample,the bitumen was preheated to approximately 105° C.-115° C. in a 10gallon feed drum and circulated through a closed feed loop system fromwhich the bitumen was fed into a semi-continuous stirred tank reactorwith vapor effluent capability, where the reactor had an internal volumecapacity of 1000 cm³. The reactor was operated in a continuous mode withrespect to the bitumen feedstream and the vapor effluent product,however, the reactor did not include a bleed stream to removeaccumulating metals and/or carbonaceous solids. The bitumen feed of eachsample was fed to the reactor as needed to maintain a working volume offeed in the reactor of 500 ml, therefore, the liquid hourly spacevelocity of the bitumen feed depended on the rate of the reaction. ABerthold single-point source nuclear level detector located outside thereactor was used to control the working volume in the reactor. 50 gramsof the catalyst was mixed with the hydrogen, hydrogen sulfide, andbitumen feed sample in the reactor during the course of thehydrocracking treatment. The bitumen feed sample, hydrogen, hydrogensulfide, and the catalyst were mixed together in the reactor by stirringwith an Autoclave Engineers MagneDrive® impeller at 1200 rpm. Vaporizedproduct exited the reactor, where a liquid product was separated fromthe vaporized product by passing the vaporized product through a highpressure separator operated at reaction pressure and 80° C. and thenthrough a low pressure separator operated at 0.17 MPa and 80° C. toseparate the liquid product from non-condensable gases.

The rate of the production of hydrocracked product was measured for eachof the samples. The results are shown in Table 2:

TABLE 2 Time [hrs] 5 10 15 20 [mol %] H₂S Rate [Kg/h · m³] 0.0% 370 335300 265 5.0% 403 370 338 305 11.4% 426 394 361 329 20.1% 448 418 387 357A graphic depiction of the rate of production of product in each of thehydrocracking reactions is shown in FIG. 3.

As shown in Table 2 and FIG. 3, the rate of production of product in thehydrocracking reactions at constant temperature and pressure increasesas the quantity of hydrogen sulfide in the reaction mixture increases.Each of the hydrocracking reactions provided a rate of at least 350kg/hr m³ for a period of time, where the rate of the reaction ismaintained above 350 kg/h-m³ for a sustained period when hydrogensulfide is present in an amount relative to hydrogen of at least 1:14where the hydrogen sulfide and hydrogen provide at least 60% of thetotal pressure in the reaction, and is sustained for a longer period asthe hydrogen sulfide levels increase.

Example 3

Another catalyst was prepared, where at least a portion of the catalysthad the structure as shown in formula (XVII) above. A 22-literround-bottom flask was charged with a solution of 1199 grams of coppersulfate (CuSO₄) in 2 liters of water. The copper sulfate solution washeated to 85° C. 520.6 grams of ammonium tetrathiomolybdate (ATTM){(NH₄)₂(MoS₄)} in 13 liters of water was injected into the heated coppersulfate solution through an injection nozzle over a period of 4 hourswhile stirring the solution. After the addition was complete, thesolution was stirred for 8 hours at 93° C. and then was allowed to cooland settle overnight.

Solids were then separated from the slurry. Separation of the slurry wasaccomplished using a centrifuge separator at 12,000 Gauss to give a redpaste. The separated solids were washed with water until conductivitymeasurements of the effluent were under 100 μSiemens at 33° C. Residualwater was then removed from the solids by vacuum distillation at 55° C.and 29 inches of Hg pressure. 409 grams of catalyst solids wererecovered. Semi-quantitative XRF (element, mass %) measured: Cu, 16.4;Mo, 35.6; S, 47.7; and less than 0.1 wt. % Fe and Co.

The catalyst solids were particulate having a particle size distributionwith a mean particle size of 47.4 μm as determined by laserdiffractometry using a Mastersizer S made by Malvern Instruments. TheBET surface area of the catalyst was measured to be 113 m²/g and thecatalyst pore volume was measured to be 0.157 cm³/g. The catalyst had apore size distribution, where the median pore size diameter wasdetermined to be 56 angstroms. X-ray diffraction and Raman IRspectroscopy confirmed that at least a portion of the catalyst had astructure in which copper, sulfur, and molybdenum were arranged as shownin Formula (XVII) above.

A solid material prepared in a similar manner was determined to have anacidity of 70 μmol of ammonia uptake per gram of solid material.

Example 4

Peace River bitumen having the composition shown in Table 1 above washydrocracked in a process in accordance with the present invention at atemperature of 454° C. Two samples of bitumen were hydrocracked underconditions the same as described in Example 2 except that the catalystused was the catalyst prepared in Example 3, the temperature of thereaction was maintained at 454° C., the pressure was maintained at atotal pressure of 13.1 MPa, and the gas flow rate of gases provided tothe reactor was 952 kg/hr m³ for one hydrocracking reaction and 949kg/hr m³ for the other hydrocracking reaction. The bitumen feed ratedepended on the rate of production of the hydrocarbon product, where thebitumen feed was fed to the reactor as needed to maintain the workingvolume of the reaction mixture. The hydrocracking conditions, bitumenfeed rate, and liquid product characteristics for the samples are shownin Table 3:

TABLE 3 Sample 1 Sample 2 Catalyst loaded (g) 50 50 Temperature (° C.)454 454 Total pressure (MPa) 13.1 13.1 Gas flow rate (SLPH) 952 949 H₂partial pressure (MPa) 8.8 8.8 H₂S partial pressure (MPa) 4.3 4.3Bitumen feed rate (g/h) 400 425 Total liquid in (kg) 17.2 17.8 Totalliquid out (kg) 14.7 14.1 Liquid recovery (wt. %) 85.2 79.0 Productdensity (g/cm³) 0.9234 0.9235 Product API Gravity (15.6° C.) 21.8 21.7Product viscosity (cSt)(15.6° C.) 10.3 10.4 Product carbon content (wt.%) 85.0 85.4 Product sulfur content (wt. %) 3.3 3.2 Product nitrogencontent (wt. %) 0.3 0.3 Boiling point fractions (wt. %-- SimulatedDistillation as per ASTM D5307) Initial boiling point - 204° C. 15.516.0 (IBP - 400° F.) 204° C.-260° C. (400° F.-500° F.) 14.5 14.5 260°C.-343° C. (500° F.-650° F.) 31.0 30.5 343° C.-538° C. (650° F.-1000°F.) 37.5 38.0 538° C.+ (1000° F.+) 1.5 1.0

Table 3 shows production of a liquid hydrocarbon product from a feedcontaining at least 20 wt. % hydrocarbons having a boiling point ofgreater than 538° C. at a feed rate of at least 400 kg/hr m³, where theliquid hydrocarbon product is recovered in a yield of at least 79% wherethe liquid hydrocarbon product contains a large proportion ofhydrocarbons having a boiling point of 538° C. or less and at most 1.5wt. % of hydrocarbons having a boiling point of greater than 538° C.

The present invention is well adapted to attain the ends and advantagesmentioned as well as those that are inherent therein. The particularembodiments disclosed above are illustrative only, as the presentinvention may be modified and practiced in different but equivalentmanners apparent to those skilled in the art having the benefit of theteachings herein. Furthermore, no limitations are intended to thedetails of construction or design herein shown, other than as describedin the claims below. It is therefore evident that the particularillustrative embodiments disclosed above may be altered or modified andall such variations are considered within the scope and spirit of thepresent invention. While compositions and methods are described in termsof “comprising,” “containing,” or “including” various components orsteps, the compositions and methods can also “consist essentially of” or“consist of” the various components and steps. Whenever a numericalrange with a lower limit and an upper limit is disclosed, any number andany included range falling within the range is specifically disclosed.In particular, every range of values (of the form, “from a to b,” or,equivalently, “from a-b”) disclosed herein is to be understood to setforth every number and range encompassed within the broader range ofvalues. Whenever a numerical range having a specific lower limit only, aspecific upper limit only, or a specific upper limit and a specificlower limit is disclosed, the range also includes any numerical value“about” the specified lower limit and/or the specified upper limit Also,the terms in the claims have their plain, ordinary meaning unlessotherwise explicitly and clearly defined by the patentee. Moreover, theindefinite articles “a” or “an”, as used in the claims, are definedherein to mean one or more than one of the element that it introduces.

1. A process for cracking a hydrocarbon-containing feedstock,comprising: continuously or intermittently providing hydrogen to amixing zone; providing a metal-containing catalyst to the mixing zone;selecting a hydrocarbon-containing feedstock containing at least 20 wt.% hydrocarbons having a boiling point of greater than 538° C. asdetermined in accordance with ASTM Method D5307; continuously orintermittently providing the hydrocarbon-containing feedstock to themixing zone at a selected rate and blending the hydrogen, thehydrocarbon-containing feedstock, and the catalyst in the mixing zone ata temperature of from 375° C. to 500° C. and at a total pressure of from6.9 MPa to 27.5 MPa to produce: a) a vapor comprised of hydrocarbonsthat are vaporizable at the temperature and the pressure within themixing zone, and b) a hydrocarbon-depleted feed residuum comprisinghydrocarbons that are liquid at the temperature and pressure within themixing zone, where the combined volume of the hydrocarbon-depleted feedresiduum, the catalyst, and the hydrocarbon-containing feedstock in themixing zone define a mixture volume in the mixing zone, wherein the rateat which the hydrocarbon-containing feedstock is provided to the mixingzone is selected to be at least 350 kg/hr m³ of the mixture volume inthe mixing zone; and separating at least a portion of the vapor from themixing zone while retaining in the mixing zone at least a portion of thehydrocarbon-depleted feed residuum and at least a portion of thecatalyst.
 2. The process of claim 1 further comprising the step of,apart from the mixing zone, condensing at least a portion of the vaporseparated from the mixing zone to produce a liquidhydrocarbon-containing product separate from the mixing zone.
 3. Theprocess of claim 2 wherein the hydrocarbon-containing product condensedfrom the vapor separated from the mixing zone contains less than 3 wt. %hydrocarbons having a boiling point of greater than 538° C. asdetermined in accordance with ASTM Method D5307 and at most 0.5 wt. %coke as determined in accordance with ASTM Method D4072, and contains atleast 80% of the atomic carbon initially contained in thehydrocarbon-containing feedstock provided to the mixing zone.
 4. Theprocess of claim 2 wherein the hydrocarbon-containing product condensedfrom the vapor separated from the mixing zone contains at least 40% ofthe atomic sulfur present in the hydrocarbon-containing feedstock. 5.The process of claim 2 wherein the hydrocarbon-containing productcondensed from the vapor separated from the mixing zone contains atleast 40% of the atomic nitrogen present in the hydrocarbon-containingfeedstock.
 6. The process of claim 2 wherein the hydrocarbon-containingproduct contains at most 0.001 wt. % vanadium and at most 0.001 wt. %nickel.
 7. The process of claim 1 wherein the vapor separated from themixing zone is hydroprocessed by contacting the vapor with ahydroprocessing catalyst and hydrogen at a temperature of from 260° C.to 425° C. and a total pressure of from 3.4 MPa to 27.5 MPa.
 8. Theprocess of claim 1 wherein the hydrocarbon-depleted feed residuum isblended with hydrogen and the catalyst in the mixing zone whileseparating at least a portion of the vapor from the mixing zone.
 9. Theprocess of claim 8 wherein the hydrocarbon-containing feedstock and thehydrocarbon-depleted feed residuum are blended with hydrogen and thecatalyst in the mixing zone while separating at least a portion of thevapor from the mixing zone.
 10. The process of claim 1 wherein thehydrocarbon-containing feedstock is continuously provided to the mixingzone, the hydrogen is continuously provided to the mixing zone, at leasta portion of the vapor is continuously separated from the mixing zone,and the hydrocarbon-containing feedstock, the hydrocarbon-depleted feedresiduum, hydrogen, and the catalyst are blended in the mixing zone fora period of at least 45 hours.
 11. The process of claim 1 furthercomprising the steps of: providing hydrogen sulfide to the mixing zoneand blending the hydrogen sulfide with the hydrocarbon-containingfeedstock, the catalyst, and hydrogen in the mixing zone, whereinhydrogen sulfide is provided to the mixing zone at a mole ratio ofhydrogen sulfide to hydrogen of at least 0.5:9.5 up to 1:1, wherehydrogen and hydrogen sulfide are provided for mixing such that thecombined hydrogen and hydrogen sulfide partial pressures provide atleast 60% of the total pressure.
 12. The process of claim 1 wherein thecatalyst comprises a metal of Column 6, 14, or of the Periodic Table ora compound of a metal of Column 6, 14, or 15 of the Periodic Table and ametal of Column(s) 3 or 7-15 of the Periodic Table or a compound of ametal of Column(s) 3 or 7-15 of the Periodic Table.
 13. The process ofclaim 1 wherein the catalyst is comprised of a material comprised of afirst metal and a second metal where the first metal comprises a metalselected from the group consisting of Cu, Ni, Co, Fe, Ag, Mn, Zn, Sn,Ru, La, Ce, Pr, Sm, Eu, Yb, Lu, Dy, Pb, Sb, and Bi, where the secondmetal comprises a metal selected from the group consisting of Mo, W, V,Sn, and Sb, where the second metal is not the same as the first metal,and wherein the material is comprised of at least three linked chainelements, the chain elements comprising a first chain element includingthe first metal and having a structure according to formula (I) and asecond chain element including the second metal and having a structureaccording to formula (II)

where M¹ is the first metal where M² is the second metal where at leastone chain element in the material is a first chain element and at leastone chain element in the material is a second chain element, and wherechain elements in the material are linked by bonds between the twosulfur atoms of a chain element and the metal of an adjacent chainelement.
 14. The process of claim 1 wherein the catalyst is comprised ofa material comprised of a first metal and a second metal where the firstmetal comprises a metal selected from the group consisting of Cu, Ni,Co, Fe, Ag, Mn, Zn, Sn, Ru, La, Ce, Pr, Sm, Eu, Yb, Lu, Dy, Pb, Sb, andBi where the second metal comprises a metal selected from the groupconsisting of Mo, W, V, Sn and Sb, where the second metal is not thesame as the first metal, and wherein at least a portion of the materialof the catalyst has a structure according to a formula selected from thegroup consisting of formula (VII), formula (IX), formula (XII), andformula (XIV):

where M is either the first metal or the second metal, and at least oneM is the first metal and at least one M is the second metal;

where M is either the first metal or the second metal, at least one M isthe first metal and at least one M is the second metal, and X isselected from the group consisting of SO₄, PO₄, oxalate (C₂O₄),acetylacetonate, acetate, citrate, tartrate, Cl, Br, I, ClO₄, and NO₃;

where M is either the first metal or the second metal, at least one M isthe first metal and at least one M is the second metal, and X isselected from the group consisting of SO₄, PO₄, oxalate (C₂O₄),acetylacetonate, acetate, citrate, tartrate, Cl, Br, I, ClO₄, and NO₃;

where M is either the first metal or the second metal, at least one M isthe first metal and at least one M is the second metal, and X isselected from the group consisting of SO₄, PO₄, oxalate (C₂O₄),acetylacetonate, acetate, citrate, tartrate, Cl, Br, I, ClO₄, and NO₃.15. The process of claim 1 wherein the catalyst is a solid particulatematerial having a particle size distribution having a median particlesize or a mean particle size of from 50 nm to 5 μm.
 16. The method ofclaim 1 wherein any catalyst provided to the mixing zone has an acidityas measured by ammonia chemisorption of at most 200 μmol ammonia pergram of catalyst.
 17. The process of claim 1 wherein thehydrocarbon-containing feedstock contains at least 30 wt. % ofhydrocarbons having a boiling point of greater than 538° C. asdetermined in accordance with ASTM Method D5307.
 18. The process ofclaim 1 wherein the hydrocarbon-containing feedstock contains at least30 wt. % of hydrocarbons that have a boiling point of 538° C. or less asdetermined in accordance with ASTM Method D5307.
 19. The process ofclaim 1 wherein the temperature in the mixing zone is selected andcontrolled to be at least 430° C.
 20. The process of claim 1 wherein thehydrocarbon-depleted feed residuum contains less than 0.02 grams ofsubstances insoluble in toluene as determined in accordance with ASTMMethod D4072, excluding the catalyst, per gram of hydrocarbon-containingfeedstock provided to the mixing zone.
 21. The process of claim 1wherein the hydrocarbon-containing feedstock is provided to the mixingzone at a flow rate of at least 400 kg/h per m³ of mixture volume. 22.The process of claim 1 wherein: the mixing zone is located in reactor;the reactor has a reactor volume; the hydrocarbon-containing feedstockand the catalyst initially provided to the mixing zone define an initialmixture volume, where the initial mixture volume is from 5% to 97% ofthe reactor volume; and where the mixture volume of the catalyst, thehydrocarbon-depleted feed residuum, and the hydrocarbon-containing feedis maintained at a level of from 10% to 1940% of the initial mixturevolume.
 23. The process of claim 22 wherein the vapor separated from themixing zone is separated from the reactor.